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FLUIDIZATION, SOLIDS HANDLING, AND PROCESSING Industrial Applications Edited by Wen-Ching Yang Siemens Westinghouse Po...

Author: Wen-Ching Yang

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FLUIDIZATION, SOLIDS HANDLING, AND PROCESSING Industrial Applications

Edited by

Wen-Ching Yang Siemens Westinghouse Power Corporation Pittsburgh, Pennsylvania

np

NOYES PUBLICATIONS Westwood, New Jersey, U.S.A.

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Copyright © 1998 by Noyes Publications No part of this book may be reproduced or utilized in any form or by any means, electronic or mechanical, including photocopying, recording or by any information storage and retrieval system, without permission in writing from the Publisher. Library of Congress Catalog Card Number: 98-18924 ISBN: 0-8155-1427-1 Printed in the United States Published in the United States of America by Noyes Publications 369 Fairview Avenue, Westwood, New Jersey 07675 10 9 8 7 6 5 4 3 2 1

Library of Congress Cataloging-in-Publication Data Fluidization, solids handling, and processing : industrial applications / edited by Wen-Ching Yang. p. cm. Includes bibliographical references and index. ISBN 0-8155-1427-1 1. Fluidization. 2. Bulk solids flow. I. Yang, Wen-ching, 1939TP156.F65F5828 1998 660 ' .284292--dc21 98-18924 CIP

PARTICLE TECHNOLOGY SERIES Series Editor: Liang-Shih Fan, Ohio State University FLUIDIZATION, SOLIDS HANDLING, AND PROCESSING: Edited by Wen-Ching Yang INSTRUMENTATION FOR FLUID-PARTICLE FLOWS: by S. L. Soo

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Contributors

John C. Chen Department of Chemical Engineering Lehigh University Bethlehem, PA

Thomas B. Jones Department of Electrical Engineering University of Rochester Rochester, NY

Bryan J. Ennis E&G Associates Nashville, TN

S.B. Reddy Karri Particulate Solid Research, Inc. Chicago, IL

Liang-Shih Fan Department of Chemical Engineering Ohio State University Columbus, OH

George E. Klinzing Department of Chemical and Petroleum Engineering University of Pittsburgh Pittsburgh, PA

Leon R. Glicksman Department of Architecture, Building Technology Program Massachusetts Institute of Technology Cambridge, MA

Ted M. Knowlton Particulate Solid Research, Inc. Chicago, IL Mooson Kwauk Institute of Chemical Metallurgy Adacemia Sinica Beijing, People’s Republic of China

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Contributors

Jack Reese Department of Chemical Engineering Ohio State University Columbus, OH

Joachim Werther Technical University HamburgHarburg Hamburg, Germany

Jens Reppenhagen Technical University HamburgHarburg Hamburg, Germany

Peter Wypych Department of Mechanical Engineering University of Wollongong Wollongong, NSW, Australia

Ellen M. Silva Department of Chemical Engineering Ohio State University Columbus, OH

Shang-Tian Yang Department of Chemical Engineering Ohio State University Columbus, OH

Gabriel I. Tardos Department of Chemical Engineering City College of City University of New York New York, NY

Wen-Ching Yang Science and Technology Center Siemens Westinghouse Power Corporation Pittsburgh, PA

Richard Turton Department of Chemical Engineering West Virginia University Morgantown, WV

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Frederick A. Zenz Process Equipment Modeling & Mfg. Co., Inc. Cold Spring, NY

Preface

This volume, Fluidization, Solids Handling, and Processing, is the first of a series of volumes on “Particle Technology” to be published by Noyes Publications with L. S. Fan of Ohio State University as the consulting editor. Particles are important products of chemical process industries spanning the basic and specialty chemicals, agricultural products, pharmaceuticals, paints, dyestuffs and pigments, cement, ceramics, and electronic materials. Solids handling and processing technologies are thus essential to the operation and competitiveness of these industries. Fluidization technology is employed not only in chemical production, it also is applied in coal gasification and combustion for power generation, mineral processing, food processing, soil washing and other related waste treatment, environmental remediation, and resource recovery processes. The FCC (Fluid Catalytic Cracking) technology commonly employed in the modern petroleum refineries is also based on the fluidization principles. There are already many books published on the subjects of fluidization, solids handling, and processing. On first thought, I was skeptical about the wisdom and necessity of one more book on these subjects. On closer examination, however, I found that some industrially important subjects were either not covered in those books or were skimpily rendered. It would be a good service to the profession and the engineering community to assemble all these topics in one volume. In this book, I have invited recognized experts in their respective areas to provide a detailed treatment

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Preface

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of those industrially important subjects. The subject areas covered in this book were selected based on two criteria: (i) the subjects are of industrial importance, and (ii) the subjects have not been covered extensively in books published to date. The chapter on fluidized bed scaleup provides a stimulating approach to scale up fluidized beds. Although the scaleup issues are by no means resolved, the discussion improves the understanding of the issues and provides reassessments of current approaches. The pressure and temperature effects and heat transfer in fluidized beds are covered in separate chapters. They provide important information to quantify the effects of pressure and temperature. The gas distributor and plenum design, critical and always neglected in other books, are discussed in detail. For some applications, the conventional fluidized beds are not necessarily the best. Special design features can usually achieve the objective cheaper and be more forgiving. Two of the non-conventional fluidized beds,recirculating fluidized beds with a draft tube and jetting fluidized beds, are introduced and their design approaches discussed. Fluidized bed coating and granulation, applied primarily in the pharmaceutical industry, is treated from the fluidization and chemical engineering point of view. Attrition, which is critical in design and operation of fluidized beds and pneumatic transport lines, is discussed in detail in a separate chapter. Fluidization with no bubbles to minimize bypassing, bubbleless fluidization, points to potential areas of application of this technology. The industrial applications of the ever-increasingly important three-phase fluidization systems are included as well. The developments in dense phase conveying and in long distance pneumatic transport with pipe branching are treated separately in two chapters. The cyclone, the most common component employed in plants handling solids and often misunderstood, is elucidated by an experienced practitioner in the industry. The book is concluded with a discussion on electrostatics and dust explosion by an electrical engineer. This book is not supposed to be all things to all engineers. The primary emphasis of the book is for industrial applications and the primary audience is expected to be the practitioners of the art of fluidization, solids handling, and processing. It will be particularly beneficial for engineers who operate or design plants where solids are handled, transported, and processed using fluidization technology. The book, however, can also be useful as a reference book for students, teachers, and managers who study particle technology, especially in the areas of application of fluidization technology and pneumatic transport.

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Preface

I’d like to take this opportunity to thank Professor Fan who showed confidence in me to take up this task and was always supportive. I’d also like to thank the authors who contributed to this book despite their busy schedules. All of them are recognized and respected experts in the areas they wrote about. The most appreciation goes to my wife, Rae, who endured many missing weekends while I worked alone in the office. Pittsburgh, Pennsylvania February, 1998

Wen-Ching Yang

NOTICE To the best of our knowledge the information in this publication is accurate; however the Publisher does not assume any responsibility or liability for the accuracy or completeness of, or consequences arising from, such information. This book is intended for informational purposes only. Mention of trade names or commercial products does not constitute endorsement or recommendation for use by the Publisher. Final determination of the suitability of any information or product for use contemplated by any user, and the manner of that use, is the sole responsibility of the user. We recommend that anyone intending to rely on any recommendation of materials or procedures mentioned in this publication should satisfy himself as to such suitability, and that he can meet all applicable safety and health standards.

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Contents

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Fluidized Bed Scale-up .............................................................. 1 Leon R. Glicksman 1.0 INTRODUCTION .................................................................................... 1 2.0 REACTOR MODELING: BED DIAMETER INFLUENCE ............... 4 3.0 INFLUENCE OF BED DIAMETER ON HYDRODYNAMICS ....... 10 3.1 Bubbling Beds................................................................................. 10 3.2 Mixing ................................................................................... 20 3.3 Influence of Bed Diameter on Circulating Fluidized Beds .......... 22 3.4 Flow Transition............................................................................... 25 4.0 EXPERIMENTAL MEANS TO ACCOUNT FOR SCALE-UP: USE OF SCALE MODELS ........................................................................... 26 4.1 Development of Scaling Parameters ............................................. 27 4.2 Governing Equations ...................................................................... 29 4.3 Fluid-Solid Forces .......................................................................... 35 4.4 Spouting and Slugging Beds .......................................................... 38 5.0 SIMPLIFIED SCALING RELATIONSHIPS ...................................... 39 5.1 Low Reynolds Number .................................................................. 39 5.2 High Reynolds Numbers ................................................................ 41 5.3 Low Slip Velocity........................................................................... 42 5.4 General Case ................................................................................... 43 5.5 Range of Validity of Simplified Scaling ....................................... 44

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Contents

6.0 FURTHER SIMPLIFICATIONS IN THE SCALING RELATIONSHIP ................................................................................... 51 6.1 Viscous Limit .................................................................................. 51 6.2 Other Derivations for Circulating Fluidized Beds ........................ 54 6.3 Deterministic Chaos ....................................................................... 55 7.0 DESIGN OF SCALE MODELS ........................................................... 56 7.1 Full Set of Scaling Relationships ................................................... 56 7.2 Design of Scale Models Using the Simplified Set of Scaling Relationships ................................................................. 61 8.0 EXPERIMENTAL VERIFICATION OF SCALING LAWS FOR BUBBLING BEDS ................................................................................ 65 8.1 Hydrodynamic Scaling of Bubbling Beds .................................... 65 8.2 Verification of Scaling Relationships for Bubbling and Slugging Beds .......................................................................... 69 8.3 Verification of Scaling Laws for Spouting Beds .......................... 75 8.4 Verification of Scaling Relationships for Pressurized Bubbling Beds ................................................................................. 76 9.0 APPLICATIONS OF SCALING TO COMMERCIAL BUBBLING FLUIDIZED BED UNITS ..................................................................... 80 10.0 HYDRODYNAMIC SCALING OF CIRCULATING BEDS ............ 91 11.0 CONCLUSIONS ................................................................................. 100 ACKNOWLEDGMENTS ........................................................................... 102 NOTATIONS ................................................................................. 103 REFERENCES ................................................................................. 104

2 Pressure and Temperature Effects in Fluid-Particle Systems .......................................................... 111 Ted M. Knowlton 1.0 INTRODUCTION ................................................................................ 111 1.1 Minimum Fluidization Velocity .................................................. 113 1.2 Bed Voidage and Bed Expansion ................................................ 120 1.3 Bubbles in Fluidized Beds ........................................................... 124 1.4 Bubble Size and Frequency ......................................................... 125 1.5 Bed-to-Surface Heat Transfer Coefficient .................................. 129 1.6 Entrainment and Transport Disengaging Height ........................ 131 1.7 Particle Attrition at Grids ............................................................. 134 1.8 Particle Attrition in Cyclones ....................................................... 136 1.9 Jet Penetration ............................................................................... 137 1.10 Regime Transitions ....................................................................... 139 1.11 Cyclone Efficiency ....................................................................... 146 NOTATIONS ................................................................................. 147 REFERENCES ................................................................................. 149

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Heat Transfer in Fluidized Beds ..........................................153 John C. Chen 1.0 INTRODUCTION ................................................................................ 153 2.0 BUBBLING DENSE FLUIDIZATION ............................................. 154 2.1 Hydrodynamic Characteristic ...................................................... 154 2.2 Heat Transfer to Submerged Surfaces ......................................... 155 3.0 CIRCULATING FAST FLUIDIZATION .......................................... 173 3.1 Hydrodynamic Characteristics ..................................................... 173 3.2 Heat Transfer ................................................................................ 178 NOTATIONS ................................................................................. 201 Subscripts ................................................................................. 202 REFERENCES ................................................................................. 202

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Gas Distributor and Plenum Design in Fluidized Beds ...209 S.B. Reddy Karri and Ted M. Knowlton 1.0 INTRODUCTION ................................................................................ 209 2.0 TYPES OF GRIDS............................................................................... 210 2.1 Perforated Plates (Upwardly-Directed Flow) ............................. 210 2.2 Bubble Cap (Laterally-Directed Flow) ....................................... 210 2.3 Sparger (Laterally or Downwardly-Directed Flow) ................... 211 2.4 Conical Grids (Laterally-Directed Flow) .................................... 211 3.0 GRID DESIGN CRITERIA ................................................................ 212 3.1 Jet Penetration ............................................................................... 212 3.2 Grid Pressure-Drop Criteria ......................................................... 214 3.3 Design Equations .......................................................................... 215 3.4 Additional Criteria for Sparger Grids .......................................... 218 3.5 Port Shrouding or Nozzle Sizing ................................................. 219 4.0 PARTICLE ATTRITION AT GRIDS ................................................ 220 4.1 Attrition Correlation ..................................................................... 222 5.0 EROSION ................................................................................. 223 6.0 EFFECTS OF TEMPERATURE AND PRESSURE ......................... 223 7.0 PLENUM DESIGN .............................................................................. 223 8.0 DESIGN EXAMPLES ......................................................................... 225 8.1 FCC Grid Design .......................................................................... 225 8.2 Polyethylene Reactor Grid Design .............................................. 230 NOTATIONS ................................................................................. 233 REFERENCES ................................................................................. 235

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Engineering and Applications of Recirculating and Jetting Fluidized Beds ...................................................236 Wen-Ching Yang 1.0 INTRODUCTION ................................................................................ 236

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Contents 2.0 RECIRCULATING FLUIDIZED BEDS WITH A DRAFT TUBE . 237 2.1 Draft Tube Operated As A Fluidized Bed .................................. 240 2.2 Draft Tube Operated As A Pneumatic Transport Tube ............. 242 2.3 Design Example for a Recirculating Fluidized Bed with a Draft Tube .......................................................................... 257 2.4 Industrial Applications ................................................................. 263 3.0 JETTING FLUIDIZED BEDS ............................................................ 264 3.1 Jet Penetration and Bubble Dynamics ......................................... 265 3.2 Gas Mixing Around the Jetting Region ...................................... 281 3.3 Solids Circulation in Jetting Fluidized Beds ............................... 295 3.4 Fines Residence Time in Jetting Fluidized Beds ........................ 315 3.5 Scale-up Considerations ............................................................... 317 3.6 Applications ................................................................................. 319 NOTATIONS ................................................................................. 319 Greek Letters ................................................................................. 322 REFERENCES ................................................................................. 323

6 Fluidized Bed Coating and Granulation ........................... 331 Richard Turton, Gabriel I. Tardos, and Bryan J. Ennis 1.0 INTRODUCTION ................................................................................ 331 2.0 COATING OF PARTICLES IN FLUIDIZED BEDS ....................... 333 2.1 Introduction ................................................................................. 333 2.2 Overview of Coating Process ....................................................... 335 2.3 Microscopic Phenomena .............................................................. 339 2.4 Modelling ................................................................................. 344 2.5 Design Criteria .............................................................................. 355 3.0 GRANULATION OF FINE POWDERS IN FLUIDIZED BEDS.... 365 3.1 Introduction ................................................................................. 365 3.2 Microscopic Phenomena .............................................................. 366 3.3 Granule Growth Kinetics ............................................................. 380 3.4 Experimental Support and Theoretical Predictions .................... 387 3.5 Granule Consolidation, Attrition and Breakage ......................... 398 3.6 Modeling of Granulation Processes ............................................ 406 3.7 Unwanted Aggregation in Fluidized Beds .................................. 418 ACKNOWLEDGMENT ............................................................................. 424 NOTATIONS ................................................................................. 424 REFERENCES ................................................................................. 429

7 Attrition in Fluidized Beds and Pneumatic Conveying Lines .................................................................... 435 Joachim Werther and Jens Reppenhagen 1.0 INTRODUCTION ................................................................................ 435 2.0 FACTORS AFFECTING ATTRITION ............................................. 437 2.1 Material Properties ....................................................................... 438 2.2 Process Conditions ....................................................................... 440

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3.0 ASSESSMENT OF ATTRITION ....................................................... 444 3.1 Breakage and Selection Functions............................................... 444 3.2 Attrition Rate................................................................................. 445 3.3 Friability Indices ........................................................................... 446 3.4 Grindability Indices ...................................................................... 446 4.0 ATTRITION TESTS ............................................................................ 447 4.1 Friability Tests .............................................................................. 447 4.2 Experiments to Study Attrition Mechanisms .............................. 448 4.3 Test Equipment and Procedures .................................................. 449 5.0 ATTRITION IN FLUIDIZED BED SYSTEMS ................................ 455 5.1 Sources of Attrition ...................................................................... 455 5.2 Attrition in the Overall Fluidized Bed System, Continuous Processes ................................................................... 473 5.3 Steps to Minimize Attrition in Fluidized Beds ........................... 475 6.0 ATTRITION IN PNEUMATIC CONVEYING LINES.................... 478 6.1 Modeling ................................................................................. 480 6.2 Parameter Effects .......................................................................... 480 6.3 Steps to Minimize Attrition in Pneumatic Conveying Lines ..... 482 NOTATIONS ................................................................................. 484 Subscripts ................................................................................. 485 Greek Symbols ................................................................................. 486 REFERENCES ................................................................................. 486

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Bubbleless Fluidization .........................................................492 Mooson Kwauk 1.0 INTRODUCTION ................................................................................ 492 2.0 FLUIDIZED LEACHING AND WASHING .................................... 492 2.1 Uniform Particles .......................................................................... 496 2.2 Mixed Particles ............................................................................. 500 2.3 Staged Fluidized Leaching (SFL) ................................................ 502 3.0 BUBBLELESS GAS/SOLID CONTACTING .................................. 502 3.1 Bubbling Fluidization and G/S Contacting Efficiency .............. 502 3.2 Species of Bubbleless G/S Contacting ........................................ 507 4.0 DILUTE RAINING FLUIDIZATION ............................................... 508 4.1 Raining Particles Heat Exchanger ............................................... 508 4.2 Experimental Verification ............................................................ 512 4.3 Baffling and Particles Distribution .............................................. 515 4.4 Pilot Plant Demonstration ............................................................ 519 5.0 FAST FLUIDIZATION ....................................................................... 523 5.1 Longitudinal Voidage Distribution ............................................. 525 5.2 Regimes for Vertical G/S Systems .............................................. 529 5.3 Radial Voidage Distribution ........................................................ 533 5.4 Modeling Fast Fluid-bed Reactors .............................................. 533 6.0 SHALLOW FLUID BEDS .................................................................. 537 6.1 Dynamics for the Distributor Zone.............................................. 537 6.2 Activated Solids Shallow Fluid Bed Heat Exchanger ................ 537

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6.3 Cocurrent Multistage Shallow Fluid Bed .................................... 541 6.4 The Co-MSFB as a Chemical Reactor ........................................ 545 7.0 FLUIDIZATION WITH NO NET FLUID FLOW ............................ 546 7.1 Levitation of Discrete Particles .................................................... 547 7.2 Semi-Fluidization through Oscillatory Flow .............................. 551 7.3 Application to Pseudo Solid-Solid Reactions ............................. 553 8.0 PARTICLES WHICH QUALIFY FOR BUBBLELESS OPERATION ................................................... 556 8.1 Powder Characterization .............................................................. 556 8.2 Improving Fluidization by Particle Size Adjustment ................. 562 9.0 WHY BUBBLING AND NOT PARTICULATE FLUIDIZATION 569 9.1 The Energy-Minimized Multiscale (EMMS) Model .................. 570 9.2 Reconciling L/S and G/S Systems ............................................... 573 10.0 EPILOGUE ................................................................................. 576 NOTATIONS ................................................................................. 576 REFERENCES ................................................................................. 578

9 Industrial Applications of Three-Phase Fluidization Systems ................................................................................... 582 Jack Reese, Ellen M. Silva, Shang-Tian Yang, and Liang-Shih Fan 1.0 INTRODUCTION ................................................................................ 582 Part I: Smelting Reduction, Paper Processing, and Chemical Processing ................................................................................. 588 2.0 SMELTING REDUCTION ................................................................. 588 2.1 Introduction ................................................................................. 588 2.2 Principles of Smelting Reduction ................................................ 590 2.3 Post-Combustion and Heat Transfer in SRF ............................... 593 2.4 Slag Layer Behavior ..................................................................... 599 2.5 Future of Smelting Reduction of Iron Ore .................................. 603 3.0 PAPER PROCESSING ........................................................................ 604 3.1 Introduction ................................................................................. 604 3.2 Chemical Pulping of Wood Chips ............................................... 605 3.3 Pulp Bleaching and Flotation De-inking ..................................... 609 4.0 CHEMICAL PROCESSING ............................................................... 614 4.1 Introduction ................................................................................. 614 4.2 Hydrotreating/Hydrocracking Petroleum Intermediates ............ 614 4.3 Fischer-Tropsch Synthesis ........................................................... 619 4.4 Methanol Synthesis ...................................................................... 621 Part II: Three-Phase Biofluidization ...................................................... 623 5.0 BIOLOGICAL APPLICATIONS OF THREE-PHASE FLUIDIZATION ................................................................................. 623 5.1 Introduction ................................................................................. 623 5.2 Applications ................................................................................. 629 5.3 Bioparticles ................................................................................. 637

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5.4 Hydrodynamics ............................................................................. 643 5.5 Phase Mixing in a Three-Phase Reactor ..................................... 647 5.6 Mass Transfer................................................................................ 648 5.7 Modeling ................................................................................. 651 5.8 Scale Up ................................................................................. 653 5.9 Process Strategy ............................................................................ 655 5.10 Novel Reactors .............................................................................. 657 5.11 Economics ................................................................................. 661 5.12 Summary ................................................................................. 662 ACKNOWLEDGMENT ............................................................................. 663 NOTATIONS ................................................................................. 663 REFERENCES ................................................................................. 664

10 Dense Phase Conveying .........................................................683 George E. Klinzing 1.0 INTRODUCTION ................................................................................ 683 2.0 ADVANTAGES OF DENSE PHASE CONVEYING ...................... 693 3.0 BASIC PHYSICS ................................................................................. 695 4.0 PULSED PISTON FLOWS ................................................................. 698 5.0 VERTICAL FLOW SYSTEMS .......................................................... 706 6.0 BOOSTERS ................................................................................. 708 NOTATIONS ................................................................................. 709 Greek ................................................................................. 709 Subscripts ................................................................................. 710 REFERENCES ................................................................................. 710

11 Design Considerations of Long-Distance Pneumatic Transport and Pipe Branching ............................................712 Peter W. Wypych 1.0 INTRODUCTION ................................................................................ 712 2.0 LONG-DISTANCE PNEUMATIC CONVEYING ........................... 713 2.1 Product Characterization and Classification ............................... 714 2.2 Blow Tank Design ........................................................................ 733 2.3 Conveying Characteristics ........................................................... 738 2.4 Pressure Drop Prediction.............................................................. 741 2.5 Stepped-Diameter Pipelines ......................................................... 747 2.6 Valves ................................................................................. 748 2.7 Pipeline Unblocking Techniques ................................................. 751 2.8 General Considerations ................................................................ 752 3.0 PIPE BRANCHING ............................................................................. 753 3.1 Dust Extraction ............................................................................. 754 3.2 Flow Splitting................................................................................ 760 3.3 Pressure Loss ................................................................................ 766 NOTATIONS ................................................................................. 767 REFERENCES ................................................................................. 769

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12 Cyclone Design ...................................................................... 773 Frederick A. Zenz 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0 9.0 10.0 11.0

INTRODUCTION ................................................................................ 773 REQUIRED DESIGN DATA ............................................................. 774 CORRELATING FRACTIONAL COLLECTION EFFICIENCY .. 775 EFFECT OF SOLIDS LOADING ...................................................... 778 CYCLONE LENGTH .......................................................................... 778 CONES, DUST HOPPERS AND EROSION .................................... 780 CYCLONE INLET AND OUTLET CONFIGURATIONS.............. 781 THE COUPLING EFFECT ................................................................. 785 PRESSURE DROP .............................................................................. 787 SPECIAL CASES ................................................................................ 788 BED PARTICLE SIZE DISTRIBUTION AND CYCLONE DESIGN ................................................................. 791 12.0 CENTRIFUGAL VERSUS CENTRIPETAL CUT POINT PARTICLE SIZE ................................................................................. 793 13.0 CYCLONE DESIGN EXAMPLES .................................................... 794 14.0 ALTERNATE APPROACH TO SOLVING EXAMPLE B ............. 804 15.0 ALTERNATE APPROACH TO SOLVING EXAMPLE C ............. 809 16.0 DIPLEG SIZING AND CYCLONE PRESSURE BALANCE ......... 812 NOTATIONS ................................................................................. 814 REFERENCES ................................................................................. 815

13 Electrostatics and Dust Explosions in Powder Handling .... 817 Thomas B. Jones 1.0 INTRODUCTION ................................................................................ 817 2.0 CHARGING OF SOLID PARTICLES .............................................. 818 2.1 Triboelectrification ....................................................................... 819 2.2 Charge Relaxation ........................................................................ 823 2.3 Induction Charging of Particles ................................................... 824 2.4 Electrostatic Fields and Potentials ............................................... 825 3.0 FLUIDIZED BED ELECTRIFICATION .......................................... 829 3.1 Background ................................................................................. 829 3.2 More Recent Work ....................................................................... 832 3.3 Beneficial Effects of Electric Charge .......................................... 836 4.0 ESD DUST IGNITION HAZARDS ................................................... 836 4.1 Basics of Suspended Solids Ignition ........................................... 837 4.2 Types of Discharges .................................................................... 841 4.3 Charge Dissipation ....................................................................... 850 5.0 ESD HAZARDS IN FLUIDIZED BED SYSTEMS ......................... 854 5.1 Hazards Associated with Fluidization ......................................... 855 5.2 Hazards in Peripheral Equipment and Processes ........................ 857 5.3 Other Nuisances and Hazards ...................................................... 863 6.0 CONCLUSION ................................................................................. 864 ACKNOWLEDGMENT ............................................................................. 866 REFERENCES ................................................................................. 867

Index .............................................................................................. 872

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1 Fluidized Bed Scale-up Leon R. Glicksman

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INTRODUCTION

Although fluidized beds have been used extensively in commercial operations such as fluidized bed combustors and fluid catalytic cracking, engineers are still faced with uncertainties when developing new commercial designs. Typically, the development process involves a laboratory bench scale unit, a larger pilot plant, and a still larger demonstration unit. Many of the important operating characteristics can change between the different size units. There is a critical problem of scale-up: how to accurately account for the performance changes with plant size to insure that a full size commercial unit will achieve satisfactory performance. In addition, it would be helpful if the smaller units could be used to optimize the commercial plant or solve existing problems. One discouraging problem is the decrease in reactor or combustor performance when a pilot plant is scaled up to a larger commercial plant. These problems can be related to poor gas flow patterns, undesirable solid mixing patterns and physical operating problems (Matsen, 1985). In the synthol CFB reactors constructed in South Africa, first scale-up from the pilot plant increased the gas throughput by a factor of 500. Shingles and McDonald (1988) describe the severe problems initially encountered and their resolution.

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Fluidization, Solids Handling, and Processing

In some scaled up fluidized bed combustors, the lower combustion zone has been divided into two separate subsections, sometimes referred to as a “pant leg” design, to provide better mixing of fuel and sorbent in a smaller effective cross section and reduce the potential maldistribution problems in the scaled up plant. Matsen (1985) pointed out a number of additional problem areas in scale-up such as consideration of particle size balances which change over time due to reaction, attrition and agglomeration. Erosion of cyclones, slide valves and other components due to abrasive particles are important design considerations for commercial units which may not be resolved in pilot plants. If mixing rates and gas-solid contacting efficiencies are kept constant between beds of different size, then thermal characteristics and chemical reaction rates should be similar. However, in general, the bed hydrodynamics will not remain similar. In some instances, the flow regime may change between small and large beds even when using the same particles, superficial gas velocity and particle circulation rate per unit area. The issue of scale-up involves an understanding of these hydrodynamic changes and how they, in turn, influence chemical and thermal conditions by variations in gas-solid contact, residence time, solid circulation and mixing and gas distribution. There are several avenues open to deal with scale-up. Numerical models have been developed based on fundamental principals. The models range from simple one-dimensional calculations to complex multidimensional computational fluid dynamics solutions. There is no doubt that such first principal models are a great aid in synthesizing test data and guiding the development of rational correlations. In a recent model evaluation, modelers where given the geometry and operating parameters for several different circulating beds and asked to predict the hydrodynamic characteristics without prior knowledge of the test results (Knowlton et al. 1995). None of the analytical or numerical models could reliably predict all of the test conditions. Few of the models could come close to predicting the correct vertical distribution of solid density in the riser and none could do it for all of the test cases! Although it is tempting to think that these problems can be solved with the “next generation of computers,” until there is general agreement and thorough verification of the fundamental equations used to describe the hydrodynamics, the numerical models will not stand alone as reliable scale-up tools. On the other hand, there is a blizzard of empirical and semiempirical correlations which exist in the fluidized bed literature to predict

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fluid dynamic behavior. In addition there are probably a large number of proprietary correlations used by individual companies. The danger lies in extrapolating these relations to new geometric configurations of the riser or inlet, to flow conditions outside the range of previous data, or to beds of much different sizes. Avidan and coauthors in a 1990 review of FCC summed up the state of the art: “basic understanding of complex fluidization phenomena is almost completely lacking. While many FCC licensors and operators have a large body of in-house proprietary data and correlations, some of these are not adequate, and fail when extrapolated beyond their data base.” (Avidan, et al., 1990.) As a example, consider the influence of mean particle size. In the early work on bubbling fluidized bed combustors, attempts were made to use relations from the classic fluidization literature which had concentrated on FCC applications with much smaller particles. In many cases, it was discovered that the relationships for small particles gave erroneous results for combustors with much larger particles. For example, the two phase theory equating the excess gas velocity above minimum fluidization to the visible bubble flow was shown to be severely distorted for large particle systems. Jones and Glicksman (1985) showed that the visible bubble flow in a bubbling bed combustor was less than one fifth of uo-umf . In other cases even the trends of the parametric behavior were changed. Heat transfer to immersed surfaces in fine particle bubbling beds increased strongly with a decrease in the mean particle size. For large particle beds, the heat transfer, in some instances, decreased with a decreased particle diameter. Another approach to scale-up is the use of simplified models with key parameters or lumped coefficients found by experiments in large beds. For example, May (1959) used a large scale cold reactor model during the scale-up of the fluid hydroforming process. When using the large cold models, one must be sure that the cold model properly simulates the hydrodynamics of the real process which operates at elevated pressure and temperature. Johnsson, Grace and Graham (1987) have shown one example of verification of a model for 2.13 m diameter industrial phthalic anhydride reactor. Several bubbling bed models gave good overall prediction of conversion and selectivity when proper reaction kinetics were used along with a good estimate of the bubble size. The results were shown to be quite sensitive to the bubble diameter. The comparison is a good check of the models but the models are incomplete without the key hydrodynamic data. In this case, the bubble size estimates were obtained from measurements of overall bed density in the reactor.

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As Matsen expresses it, after over a half a century of scale-up activity in the chemical process industry, “such scale-up is still not an exact science but is rather a mix of physics, mathematics, witchcraft, history and common sense which we call engineering.” (Matsen, 1995.) A complete treatment of scale-up should include the models, numerical calculation procedures and experimental data designers need to carry out successful scale-up from small size beds to commercial units. This would involve a large measure of the existing fluidized bed research and development effort; clearly, such a task is beyond the scope of a single chapter. Since changes in the bed size primarily influence scale-up through changes in the bed hydrodynamics, one focus of this chapter is on experimental results and models which deal explicitly with the influence of bed diameter on hydrodynamic performance for both bubbling and circulating fluidized beds. The changes in the bed dynamics will, in turn, impact the overall chemical conversion or combustion efficiency through changes in the particle-to-gas mass transfer and the heat transfer from the bed to immersed surfaces or the bed wall. Several examples of this influence are also reviewed. The second focus of this chapter is on the use of small scale experimental models which permit the direct simulation of the hydrodynamics of a hot, possibly pressurized, pilot plant or commercial bed. By use of this modeling technique, beds of different diameters, as well as different geometries and operating conditions, can be simulated in the laboratory. To date, this technique has been successfully applied to fluidized bed combustors and gasifiers. Derivation of the scale modeling rules is presented for a variety of situations for gas solid fluidized beds. Verification experiments and comparisons to large scale commercial systems are shown. Rules for the use of this experimental modeling technique for FCC operations as well as for the simulation of bed-to-solid surface heat transfer are also given.

2.0

REACTOR MODELING: BED DIAMETER INFLUENCE

In this section, representative results are reviewed to provide a prospective of reactor modeling techniques which deal with bed size. There probably is additional unpublished proprietary material in this area. Early studies of fluidized reactors recognized the influence of bed diameter on conversion due to less efficient gas-solid contacting. Experimental studies were used to predict reactor performance. Frye et al. (1958) used

24-Sep-2001 JMR


5

a substitute reaction of ozone decomposition to study hydrocarbon synthesis. The ozone decomposition can be run at low pressures and temperatures and can be rate-controlled in the same way and by the same catalyst as the reaction under development. Frye and coworkers used three beds of 2 inch, 8 inch and 30 inch diameter, respectively, to study the size influence. We should interject a caution that the use of pressures and temperatures different than the actual reaction may mean that the hydrodynamics of the substitute reaction model will differ from the actual application; this is illustrated later in the chapter. Figure 1 shows the apparent reaction rate constant for the different bed diameters at two different bed heights with the other parameters held constant. Note that the rate constant decreased by roughly a factor of three between the 2 inch and 30 inch beds. ,)l'O.. Bx10-s

6x10-.

4x10-o

~/AL LIN'.A. ..'.l«lrr ..A,"ro. '..rtr M'.SS. .r,..~. AN~ INL'.rl

...: z ~ tn

CArALrSr ~.rl,"L'.

.'1'.

D..J ,~ "-1:

,7: " .JI:C. I: ..I ',I'. I', , ..D , I:ACII DIA. 01" .,..,,1#MJ-D

2x10-

Z O U 1-

~

IxI0-.

Z

Bx10-.

u

6xI0-.

0 i=

Ir

1- 4xI 0-&

11.

2x10..

1x10:-8

Figure 1. Apparent reaction-rate constant vs. reactor diameter and bed height. (From Frye et al., 1958.)

9-Oct-2001 JMR

6


May (1959) reports results of tests done in cold models used to simulate the flow through large reactors whose performance had been found to be inferior to that of smaller pilot units. The importance of this problem can be appreciated from the scale of the equipment used. Figure 2 shows the 5 foot diameter unit used for the scale-up tests. This unit was fluidized with compressed air at 27 to 38°C (80 to 100°F) and pressures up to 689 KPa (100 psi). Gas residence time in the bed was determined by the use of tracer gas. Radioactive solid tracers were introduced into the bed to determine solid mixing. The data obtained in the larger units are much more erratic with evidence of large scale mixing patterns. Figure 3 shows the axial mixing coefficients obtained in experiments with different size beds. Mixing in the larger diameter bed is an order of magnitude larger than that in a small laboratory unit. The measured hydrodynamic behavior of the gas and solid was combined with a reaction model to predict the reactor behavior. Here again, there should be concern about the accuracy with the air experiments done at ambient temperature. Use of identical bed geometry and bed solid material does not guarantee identical hydrodynamics. The shift in gas properties from the cold model to the hot reactor may cause a marked difference in behavior. Additional scaling parameters must be maintained constant between the reactor and the cold model to insure identical hydrodynamics, and in some cases just to guarantee identical flow regimes! Volk et al. (1962) show the effect of bed diameter on the conversion of CO in the “Hydrocol” reaction in which hydrogen and carbon dioxide are converted over an iron catalyst to hydrocarbons and oxygenated hydrocarbons in a bubbling or possibly slugging bed. Figure 4 shows the CO conversion. It is seen that the conversion rate is reduced as the reactor diameter increases. Volk used vertical tubes within the reactor to reduce the equivalent diameter of the system, equal to the hydraulic diameter, four times the free cross sectional area divided by the wetted perimeters of all surfaces in the cross section. The performance was found to be correlated by the equivalent diameter. It was also found that bed expansion was correlated with bed diameter. In their process, larger beds were built with internals which kept the equivalent diameter the same as that of smaller units. The large units with internals appeared to give comparable gas-to-solid contacting. The use of vertical internals may not be feasible for a number of reasons, such as tube erosion. The use of the equivalent diameter approach may not be universally valid.

24-Sep-2001 JMR


7

Figure 2. Very large equipment built to study scale-up problems. (From May, 1959.)


8

Figure 3. Solid diffusivity in axial direction for large units. (From May, 1959.)

1001

~~in::

'

12 'r'!. c o "in ~

~

c o u

o u

5 in.

i

O

801 Key

70

Symbol Description O Open'Reactor .Modified Internals , 2

I 4

16 EQuivalent

Figure

4. CO conversion

(From Volk et al., 1962.)

in Hydrocol

reaction

10 Diameter,

12

14

in.

for several reactor diameters.


9

Van Swaaij and Zuiderweg (1972) used the ozone decomposition reaction to study the conversion characteristics in a bubbling bed. Studies were made with beds of 5, 10, 23, 30 and 60 cm diameter and up to 300 cm bed heights. The results were compared with predictions using a twophase flow model with the mass transfer coefficient between the bubble and dense phase derived from residence time distribution results of gastracer pulse response tests. Figure 5 shows the height of the mass transfer unit Hα , which is equivalent to u/α where α is the mass transfer coefficient, as a function of the bed diameter. The results from the ozone conversion and the residence time distribution interpreted by the two phase model gave reasonably similar results. In these cases, the mass transfer between phases is the limiting resistance for the reaction. Note that for larger bed diameters the mass transfer coefficient decreases. Van Swaaij and Zuiderweg (1973) showed that the inclusion of vertical tubes in a bed gave bubble to dense phase mass transfer results which were roughly equivalent to a smaller open bed with the same hydraulic diameter while the solids axial mixing was higher than that predicted using the hydraulic diameter. 6 .FROM .O

CONVERSION

C o

FR)M

( PRESENT

RTD

DATA TESTS

INVESTIGATION

)

u «{H

O.i

I

4

.

.

1

I

4 D,m

Figure 5. Mass transfer unit for ozone conversion for different bed diameters. (From Van Swaaij and Zuiderweg, 1972.)

9-Oct-2001 JMR

10


Bauer et al. (1981) measured the influence of bed diameter on the catalytic decomposition of ozone. Figure 6 shows the decrease of the conversion with bed diameter for Bauer’s data. This figure also shows the influence of distributor design on conversion. In many small scale experiments, a porous plate is used which will give better performance than the distributors used in large shallow bed commercial designs. Avidan and Edwards (1986) successfully scaled up from bench scale to demonstration plant from 0.04 m to 0.6 m diameter while maintaining nearly 100% conversion for a fluid bed methanol to gasoline process. In this case, they ran at a superficial gas velocity which was high enough to be in the turbulent flow regime suppressing bubbles. By this technique they eliminated the losses associated with gas bypassing in bubbles.

Figure 6. Conversion catalytic decomposition of ozone for different bed diameters and distributors. (From Werther, 1992.)

3.0

INFLUENCE OF BED DIAMETER ON HYDRODYNAMICS

3.1

Bubbling Beds

In the studies mentioned above, the major objective was the experimental determination of conversion as a function of overall design

9-Oct-2001 JMR


11

parameters and particle properties. There have also been studies which have looked at the influence of bed diameter on the hydrodynamics in an effort to understand the cause of the conversion loss with bed size increase. These studies have aided in the development of physical models of reactor performance. De Groot (1967) measured gas residence time, bed expansion and solid axial mixing in a series of beds at different diameter fluidized with air at ambient conditions. He used a narrow size range and broad size range of crushed silica with sizes below 250 µm. Beds with diameters of 0.1, 0.15, 0.3, 0.6, and 1.5 m were used in the tests. There was a substantial decrease in bed expansion and bubble fraction for narrow size range particles at large bed diameters, indicating the possibility of gas bypassing in bubble channels (Fig. 7). The axial diffusivity also increased with bed diameter and was a strong function of particle size distribution (Fig. 8). % BED

EXPANSION

T

40 BROAO-RANGE

1--

30

.

T

.!0

'~---1---

SILICA T

---+---

..1

20 NAllROW-RANGE

SILICA

10

0 0.1

0.5

t.O

5 BED

D:AMETER,

m

9-Oct-2001 JMR

12


Figure DeGroot,

8.

Effective 1967.)

axial diffusivity

coefficients

for solids mixing.

(From


13

Weather (1974) measured the bubble characteristics in cylindrical beds of diameters 100, 200, 450 and 1000 mm, respectively, for fine particles with a mean particle diameter of 83 µm. He showed that for beds 200 mm or smaller, common sizes used for laboratory experiments, the bed diameter had a strong effect on the bed hydrodynamics. There was a zone of preferred bubble flow near the bed walls at lower elevations (Fig. 9). The bubbles grew in size and moved toward the centerline, presumably by coalescence, higher up in the bed (Fig. 10). The transition to slugging occurred higher up in the larger bed at the same superficial gas velocity. The bubble velocity increased with height until slug flow was approached, after which the velocity decreases. For the only case shown, the 100 mm bed, the maximum velocity occurred when the bubble dimension was about one third of the bed diameter. In larger beds, the bubble rise velocity was higher for the same bubble volume (Fig. 11). Hovmand and Davidson (1971) reviewed data on bubble rise velocity and concluded that the bubble rise velocity is governed by the bed diameter rather than the bubble diameter when the bubble diameter exceeded 1/3 of the bed diameter. Note that Weather’s results at different superficial velocity are well correlated by the drift flux form,

u b ~ Φ( u o - u mf ) + c gd v

Eq. (1)

:76,

I -NI

... E 721 I.J. .~ ~

~

I\

,

.0: '50 A; 1000

\

I i

,

~ \

.2 ...

In & 01 ~ .C) '1 :) .C) O {'1 u .2

1 70

20 distance

I.""'~

30 from

1.0 column

centre.

50 r.

cm

Figure 9. Bubble gas flow Vb as a function of the distance r from the vessel center line in a height of 30 cm above the distributor in beds of different diameters DB (uo = 9 cm/sec, H = 50 cm). (From Werther, 1974.)

9-Oct-2001 JMR

14


,.. > 1

where u´ = u/u o and v´ = v/uo and u´- v´ is a dimensionless slip velocity. The minimum fluidization velocity can be expressed as,

Eq. (47)

(

) (

∆p = ρ s - ρ f g 1 - ε mf L

)

(

1.75 1 - ε mf

~

ε 3mf

)

ρf

u 2mf φdp

rearranging and using ρs in place of ρs - ρf ,

Eq. (48)

(

1.75 1 - ε mf ε 3mf

)

ρf ρs φ d p

=

(

g 1 - ε mf

)

u 2mf

Substituting this into Eq. (46) and multiplying by Fr, Eq. (49) Fr

3 2 ε 3 L | u ′ - v ′ | g (1 - ε ) βL mf u 2o ε mf (1 - ε ) | u ′ - v′ | = = uo ρs u o gL ε u 2mf ε u 2mf

At large particle Reynolds numbers, just as at low Reynolds numbers, the dimensionless drag, βL/ρs uo , is identical when uo /umf , εmf

9-Oct-2001 JMR

42


and Fr are identical. The variables ε, u´ and v´ are dependent dimensionless variables which are identical for two similar fluidized beds. In this limit the same set of governing dimensionless parameters applies as in the low Reynolds number limit, given by Eq. (45). 5.3

Low Slip Velocity

Finally, consider the case when the magnitude of the slip velocity between the particles and the gas is close to umf /ε everywhere in the fluidized bed. With the vertical pressure drop equal to the particle weight, the following holds for any value of the particle Reynolds number,

Eq. (50)

∆p u mf = β | u - v |≈ β = ρs g (1 - ε ) L ε

Eq. (51)

β = ρs g (1 - ε )

ε u mf

and

Eq. (52)

1 u mf βL Fr = u o / u mf (1 - ε )ε for u′ - v′ → ρs uo ε uo

Again, when uo /umf and Fr are identical for two beds and the slip velocity is close to umf /ε, the dimensionless drag coefficient is also identical for two beds. For all three limiting cases identified above, similitude can be obtained by maintaining constant values for the dimensionless parameters,

Eq. (53)

9-Oct-2001 JMR

u o2 ρs uo L Gs , , , 1, , bed geometry, φ s , PSD gL ρ f u mf L 2 ρs u o

Fluidized Bed Scale-up 5.4

43

General Case

Since the same simplified set of dimensionless parameters holds exactly at both high and low Reynolds numbers, it is reasonable to expect that it will hold, at least approximately, over the entire range of conditions for which the drag coefficient can be determined by the Ergun equation or an equation of similar form. The validity of the simplified parameters can be checked numerically for the intermediate range of values. From Eq. (31), Eq. (54) βL ρs u o

=

150(1 - ε )2 ε

2

µL

ρs uo (φ d p )

2

+

1.75(1 - ε ) ρ f ρs ε

L | u ′ - v′ |

φ dp

where u´- v´ can also be represented as uslip /uo. Using the Ergun equation to determine umf , Eq. (55)

(1 - ε mf ) ∆p = ρs g (1 - ε mf ) = 150 L ε 3mf

2

µ umf

(φ d p )2

+

1.75(1 - ε mf ) ρ f u 2mf ε 3mf

φ dp

Dividing Eq. (54) by 55, and rearranging,

 β L  Fr (1 - ε mf )ε    ρ u  o  s   u o (1 - ε )2 3 ε mf  u mf    2

Eq. (56)

  1.75 φ s Re ε | u ′ - v′ | 1 + 150 (1 - ε )  =   1.75 φ s Re u mf   1+  150 (1 - ε mf ) u o  

where Re = ρf uo dp /µ .

9-Oct-2001 JMR

44


It is easy to verify the three limits defined previously by use of Eq. (56). For the more general case, Fig. 21 shows the value of β given by Eq. (56) relative to β at low Re over a range of conditions when uo /umf is 10 and 3, respectively, and Fr and φs remain constant. When uo /umf and the slip velocity are high, there is a larger variation of dimensionless drag coefficient with Reynolds number. Note that β does not vary with particle Reynolds number when the Reynolds number remains above about 103 or below about 10. Figure 22 illustrates the results when uo /umf is 1000, a condition approached with very fine particle bubbling beds or circulating beds. In the latter instance, the use of the Ergun relationship is questionable except for the dense lower part of the bed. 5.5

Range of Validity of Simplified Scaling

To determine the validity of the simplified scaling laws over a wide range of conditions the simplified scaling laws have been used (Eq. 53) to design hypothetical models whose linear dimensions are 1/4 and 1/ 16, respectively, of the linear dimensions of a model designed using the full set of scaling laws, Eq. (37). To determine the validity of the smaller, simplified models, the dimensionless drag coefficient β L/ρs uo will be compared between the simplified models and the model using the full set of scaling laws. Figure 23 shows a comparison of the exact model and the simplified models for a pressurized fluidized combustor. Using the full set of scaling laws, the exact model, fluidized by ambient air, is approximately the same size as the combustor. The simplified models are reduced in size by their respective assumed length scale. The other parameters of the simplified model are then calculated to match the simplified parameters. For example, when the length scale is reduced to 1/4 that of the exact model, the velocity is reduced by 1/2 to keep the Froude number constant. The particle diameter is then reduced appropriately to keep the ratio of uo /umf constant. These calculations were carried out over a range of particle Reynolds numbers, RepE, based on the full scaling law, or exact, model. In the simplified scaling relationships, the Reynolds number is not maintained constant. The concern is how much the drag coefficient is impacted by the shift in Reynolds number. It was found that the particle Reynolds number for the 1/4 scale simplified model remained roughly equal to 0.34 RepE over a wide range of values for RepE , whereas the particle Reynolds number for the 1/16 scale model was roughly 0.12 RepE.

9-Oct-2001 JMR


45

10

6' " . ~

!

1 ,

,0

R. .pUoda'1i 100

1000

10000

Uo/Umf= 10 and 3. (From Glicksman et al., 1993b.) Figure 21. Ratio of dragcoefficient to low Reynoldsnumber drag coefficient,

10

UsliplUo

= 11200

Re

100 = pUOdp'J!

O II Q) a: ~ ~

1 1

10

1000

10000

Figure 22. Ratio of drag coefficient to low Reynolds number drag coefficients, uo /umf = 1000. (From Glicksman et al., 1993b.)

9-Oct-2001 JMR

46


114ScaM SI.~n.4 M...

rn

tI 1, Scai8 Si8p6~ .~... 0!

Figure 23. Exact and simplified models of a pressurized fluidized bed combustor. (From Glicksman et al., 1993b.)

Using these Reynolds number scale factors, the errors in the dimensionless drag coefficient β L/ρs uo using the simplified scaling models are shown in Figs. 24 and 25 for uo /umf of 10 and 1000, respectively, plotted as a function of RepE based on parameters for the exact scaled bed. For a particle Reynolds number of 1000 or less, which corresponds to

9-Oct-2001 JMR


47

pressurized beds with particles of 1 mm or less, the error in the drag coefficient with the simplified scaling laws is twenty percent or less for a one quarter length scale model. The error is forty percent or less for a one sixteenth length scale model. At uo /umf of 1000 and uslip /umf of 1/50, the errors for the 1/16 scale model are twenty percent or less for RepE less than 103. For particles of 0.2 mm or less, corresponding to a Reynolds number of 100 or less, the errors in drag coefficient are minimal. When the Ergun equation applies for the drag coefficient, a one quarter scale model based on the simplified scaling laws should be valid for any conditions. A 1/16 scale model should be valid for particle diameters of about 0.2 mm or less for a pressurized bubbling bed with uo /umf of 10 and uslip /Uo of 0.3. At uo / umf of 1000 and uslip /Uo of 1/50, the 1/16 scale model should be valid for pressurized beds with particles up to 1 mm in diameter. These conclusions apply when the particle to fluid drag term is given by the Ergun equation or similar relationships and the scaled particles are not so small that interparticle surface forces come into play. Clusters. In the freeboard of a bubbling bed or in the upper portion of a circulating bed where particles generally are considered to act in clusters or groups, a similar examination of scaling of the gas-to-solid drag can be made. Consider all of the particles grouped into clusters with an effective diameter dc and the clusters occupying a volume fraction εc of the bed volume. The cluster-to-gas drag will be represented by the drag coefficient for a solid sphere of diameter dc ,

Eq. (57)

β| u - v |

π d 3c 1 π d 2c = | u - v|2 C D ρf ( ) 6 1- εc 2 4

This can be rewritten as,

Eq. (58)

βL 3  ρ f = ρs u o 4  ρs

  | u ′ - v ′ | C D L (1 - ε c)  dc 

If the diameter of individual particles does not influence the drag of a cluster of particles, then when the solid-to-gas density ratio is held constant between the combustor and the model, the dimensionless drag β L/ρs uo is properly scaled when CD is invariant.

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48


..

..

ULexact

0.3

,

a 1118

10

100 Re-

1000

1 0000

Uo~

Figure 24. Error in drag coefficient for simplified scaling relationships, uo /umf = 10. (From Glicksman et al., 1993b.)

,--':;0"

..

..

ULex.ad

a 1118

0.3 1

10

1000

100

A.-

10000

Uod,

Figure 25. Error in drag coefficient for simplified scaling relationships, uo /umf = 1000. (From Glicksman et al., 1993b.)

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49

If the reduced scale models faithfully reproduce the dynamics of the exact case, the cluster dimensions should scale directly with the linear dimensions of the bed. Thus, a one-quarter linear scale model which has a velocity one-half that of the exact case will have a cluster Reynolds number ( Re d c ) one-eighth that of the exact bed. From the relationship of CD with Re, the change of CD with model scale at a given Reynolds number of the exact bed can be determined. Figure 26 shows the shift in CD using the CD relationship of White (1974) for length scales of 1/4, 1/8 and 1/16, respectively, of the exact bed length as a function of the cluster Reynolds number of the exact bed. Also shown on the figure is the typical Reynolds number of an atmospheric combustor with a 0.3 m cluster diameter, approximately 1.5 × 104. In a bubbling bed, the cluster diameter in the freeboard should be at least equal in size to the diameter of bubbles erupting at the bed surface. For beds with horizontal tubes, the bubble diameter will be equal to or larger than the horizontal tube spacing. In a bubbling bed without tubes, the bubbles and clusters can be much larger. In an open circulating bed the cluster diameter is more difficult to determine. It is reasonable to assume its diameter is proportional to the bed diameter, equal in magnitude to the bed diameter or one order of magnitude smaller. From these considerations, the Reynolds number based on the cluster diameter should be 104 or larger in an atmospheric combustor with a cluster diameter of 0.2 m. The cluster Reynolds number should be 105 or larger in a pressurized combustor. From Fig. 26, it can be seen that a one quarter scale or an eighth scale model should have drag coefficients similar to the exact bed. For pressurized beds, the drag coefficients should be very close in magnitude. Individual Particles. If the drag coefficient, β, is influenced by the characteristics of individual particles, the detailed particle dynamics of the simplified scale models must be examined. In this case

Eq. (59)

βL 3  ρf =  4  ρs ρs u o

  | u ′ - v′ | C D L (1 - ε )  dp 

where CD is the drag coefficient of a single particle. This can be rewritten in terms of the single particle terminal velocity which can be found from

Eq. (60)

ρs π d p3 g 6

2 1 ρf π dp = CD u t2 2 4

9-Oct-2001 JMR

50


10

~ x O ~ 8

1 1000

10000

100000

Figure 26. Error in cluster drag coefficient for fixed uo /umf using Cd for a solid sphere. (From Glicksman et al., 1993b.)

Substituting Eq. (60) into Eq. (59) to eliminate CD, one obtains

βL u2 Fr = o2 | u ′ - v′ | (1 - ε ) ρs uo ut

Eq. (61)

Since uo /umf and Fr are held constant in the simplified scaling process, we will examine the ratio ut /umf to determine if the drag coefficient β L/ρs uo remains constant. The Ergun equation can be solved to find umf Eq. (62)

(

- (150) 1 - ε mf ρf

u mf d p µ

3 ε mf

=

)

2

(

 - (150) 1 - ε mf +  3  ε mf

(

)

2

)

3.5 1 - ε mf φ S 3 ε mf

where Ar = ρs ρ f d 3p g µ 2 is the Archimedes number.

9-Oct-2001 JMR

2

(

 7 1 - ε mf  + 3  ε mf

)

2 3

φ s Ar


51

The relationship given by White (1974) for CD written as,

Eq. (63)

4 Ar 3 2 Ret = 24 6 + + 0.4 Ret 1+ Ret

where Ret = ρ f ut d p µ can be used to relate ut to Ar. At small and large values of Ar, the ratio of ut/umf which is the ratio of Ret ( ρf u mf d p µ ) approaches a constant value. In these two limits, the simplified scaling laws will yield exact agreement of ut /umf between the combustor and the simplified models. The errors in ut /umf are shown in Figs. 27 and 28 for simplified scale models at two different linear dimensions. Scaling a combustor with comparatively small particles, 0.2 mm or less, gives good agreement for ut /umf even at one-sixteenth linear scale, while for large particles a linear scale of one-fourth gives fair agreement for ut /umf. Since uo /umf is held constant in the simplified scaling laws, close agreement of ut /umf also results in close agreement of ut /uo . For a circulating bed, pneumatic transport, or other low solids volume fraction, it might be more correct to scale the model to achieve closer agreement for ut /uo , allowing uo /umf to deviate between the model and full scale bed.

6.0

FURTHER SIMPLIFICATIONS IN THE SCALING RELATIONSHIP

6.1

Viscous Limit

Glicksman (1984) showed that the list of controlling dimensionless parameters could be reduced if the fluid-particle drag is primarily viscous or primarily inertial. The standard viscous and inertial limits for the drag coefficient apply. This gives approximately Viscous : Re urel ,dp | u ′ - v ′ | < 10

Eq. (64) Inertial : Reu rel ,dp | u ′ - v′ | > 1000

using the Ergun equation.

9-Oct-2001 JMR

52


1116

ULexact

= 114

1.1

, 0.1

1 dp (combustor). mm

Figure 27. Error in terminal velocity using simplified scaling at 1013 kPa and 800°C. (From Glicksman et al., 1994.)

dp (combustor)(mm)

Figure 28. Error in terminal velocity using simplified scaling at 101.3 kPa and 800°C. (From Glicksman et al., 1994.)

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53

The use of a Reynolds number based on relative velocity rather than superficial velocity in setting these limits was suggested by Horio (1990). In setting viscous or inertial limits, it is the interphase drag which is characterized as being dominated by viscous or inertial forces. The particle inertia is important even if the interphase drag is viscous dominated. This is because of the typically large solid-to-gas density ratio. For viscous dominated flows, it can be assumed that the gas inertia and the gas gravitational forces are negligible. By dropping the gas inertia and gravity time from the gas momentum equation and simplifying the dimensionless drag coefficient to the linear viscous term, the set of dimensionless equations does not include gas-to-solid density ratio as a parameter. The ratio between the bed and particle diameters and the Reynolds number based on bed diameter, superficial velocity, and solid density appear only in the modified drag expression, in which they are combined, see Eq. (40). These parameters form a single parameter, as discussed by Glicksman (1988) and other investigators. The set of independent parameters controlling viscous dominated flow are then Eq. (65)

ρs u o d 2p µL

gL

, φs ,

u 2o

,

D Gs , , bed geometry, PSD L ρs u o

The first term in the list multiplied by the third term has been shown by Glicksman (1988) to be equivalent to the ratio of superficial and minimum fluidization velocities in the viscous limit. The controlling parameters can therefore be written as

Eq. (66)

u mf uo

, φs,

gL u2o

,

D , Gs , bed geometry, PSD L ρs uo

Note this is a subset of the simplified scaling laws presented above with the solid-to-gas density ratio removed. The viscous limit requirement is based on the assumptions that the gas inertial forces are negligible compared to gas-solid interaction forces and the dimensionless interphase drag can be modelled as a linear function of the particle Reynolds number. Because of clustering of particles, this may not be a valid assumption, and the requirements for neglecting gas inertia may be much more stringent than given by Eq. (64). As will be discussed in the experimental verification of simplified scaling laws, a viscous limit for interphase drag may not exist in circulating fluidized beds.

9-Oct-2001 JMR

54


It must be borne in mind that this set is valid only when fluid inertial effects are negligible, i.e., it is a subset of the general relationships and the simplified relationship. Glicksman (1984) used the criteria for the viscous limit in a bubbling bed that the ratio of viscous forces to fluid inertial forces in the Ergun equation is ten or larger. From Table 1 of that reference, for a bed of glass or sand fluidized with air at standard conditions with uo /umf of 3, the viscous limit occurs when particles are less than about 200 microns. In regions where particles behave individually, the viscous limit occurs for particles less than 60 microns (assuming the particle/gas relative velocity is equal to the particle terminal velocity and the single particle viscous limit criterion is Reurel , dp < 1 ). There are some data to suggest this Reynolds number limit may be too restrictive. 6.2

Other Derivations for Circulating Fluidized Beds

The most dramatic departure from the above procedure for derivation of scaling laws is that of Horio (1989). Although the procedure is different, the results are identical to the scaling laws described in the previous sections. Horio has presented a scaling law for circulating fluidized beds which is based on the Clustering Annular Flow Model (CAFM) of Ishii et al. (1989). The CAFM describes the flow in a circulating fluidized bed as particle clusters moving upward in the core and downward in an annulus at the wall. The analysis was similar to that of Nakamura and Capes (1973). Completeness of the model was achieved by assuming minimization of pressure gradient. Horio’s scaling law derivation was based on the requirement that two similar circulating fluidized beds have equal values of voidage distribution, dimensionless core radius, gas splitting to core and annulus, solid splitting to core and annulus, and cluster voidage. The CAFM equations were then examined to determine how these requirements could be met. Questions as to the validity of the theory supporting Horio’s development are based on deficiencies of the CAFM to fully describe circulating fluidized bed flow and the lack of justification for the premise that the five characteristics comprise all that is of interest in a circulating fluidized bed. In particular, the assumption of pressure drop minimization, although convenient, is not justified physically. The scaling law proposed by Horio for circulating fluidized beds can be shown to be equivalent to the simplified set of parameters. Horio also discussed reductions to his list of scaling parameters in which the solid/gas density ratio is omitted. He claimed that this reduced scaling law

9-Oct-2001 JMR


55

is sufficient for “macroscopic similarity” and sacrifices only cluster size similarity. This reduced law is identical to the bubbling bed scaling law developed by Horio (1986) and has been shown by Glicksman (1988) to be equivalent to the viscous limit scaling law presented in the previous section. Horio’s suggestion that the density ratio can be neglected because sacrificing cluster size does not necessarily alter the macroscopic behavior of a circulating fluidized bed is somewhat curious since his development is based on the CAFM which assumes clusters are the primary mechanism of particle transport. From an operational standpoint, the paramount question for a designer is which scaling law must be followed. Since Horio’s reduced set is equivalent to the viscous limit, Eq. (66), it must be determined if the density ratio can be omitted in the scaling relationship, i.e., is the viscous limit an adequate set of scaling parameters or must the density ratio also be included as is done in the simplified scaling relationships, Eq. (53). As will be shown in a later section, density ratio is not a parameter which can be omitted when scaling circulating fluidized beds and the density ratio is generally required for scaling bubbling beds. 6.3

Deterministic Chaos

Recent studies have indicated that fluidized beds may be deterministic chaotic systems (Daw et al.,1990; Daw and Harlow, 1991; Schouten and van den Bleek, 1991; van den Bleek and Schouten, 1993). Such systems are characterized by a limited ability to predict their evolution with time. If fluidized beds are deterministic chaotic systems, the scaling laws should reflect the restricted predictability associated with such systems. Van den Bleek and Schouten (1993) have suggested that if two beds are properly scaled, the rate of information change in both systems will be the same. They suggest that two scaled beds will exhibit the same Kolmogorov entropy, or information generation rate, when measured on the same time scale. Hence, they adopt an additional method for verifying dynamic similarity requiring that the information group, Kt, remain constant (where t = dp /uo and K is the Kolmogorov entropy) along with the full set of scaling parameters we have described. Further work is needed to determine in which regimes, if any, fluid bed behave as chaotic systems. Additional testing is needed to determine the sensitivity of important bed hydrodynamic characteristics to the Kolmogorov entropy, to quantitatively relate changes of entropy to

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56


shifts in hydrodynamic behavior. It may be that the requirement of Kt equality may be an over-restrictive condition, that is, similarity of hydrodynamic characteristics such as solid mixing and gas-to-solid transfer may be achieved even when Kt varies considerably.

7.0

DESIGN OF SCALE MODELS

7.1

Full Set of Scaling Relationships

We will first consider the steps to design a model which is similar to another bed based on the full set of scaling parameters, Eqs. (37) or (39). To construct a model which will give behavior similar to another bed, for example, a commercial bed, all of the dimensionless parameters listed in Eqs. (37) or (39) must have the same value for the two beds. The requirements of similar bed geometry is met by use of geometrically similar beds; the ratio of all linear bed dimensions to a reference dimension such as the bed diameter must be the same for the model and the commercial bed. This includes the dimensions of the bed internals. The dimensions of elements external to the bed such as the particle return loop do not have to be matched as long as the return loop is designed to provide the proper external solids flow rate and size distribution and solid or gas flow fluctuations in the return loop do not influence the riser behavior (Rhodes and Laussman, 1992). Proper conditions must be chosen to design a scale model to match the dimensionless parameters of the commercial bed. To model a gas fluidized commercial bed, a scale model using air at standard conditions is most convenient, although several investigators have used other gases (Fitzgerald and Crane, 1980; Fitzgerald et al., 1984; Chang and Louge, 1992) or pressurized scale models (Almstedt and Zakkay, 1990; Di Felice et al., 1992 a,b). The gas chosen for the model, along with the gas pressure and temperature, determines the values of ρf and µ. The particle density for the model is chosen to match the density ratio, so that

Eq. (67)

 ρf  ρ  s

 ρ   =  f     m  ρs c

where the subscript m is for the model and c is for the commercial bed. For the remaining parameters, the form of Eq. (37) will be chosen for the

9-Oct-2001 JMR


57

dimensionless parameters. Combining the Reynolds number based on bed diameter and the square root of the Froude number,

Eq. (68)

 32 g   32 g  ρ f uo L gL L    = = L  νf   νf  µf uo  m  c

Rearranging, 2

Eq. (69)

 Lm  =  (ν f )m  3   ( )   Lc   ν f c 

All of the linear dimensions of the model are scaled to the corresponding dimensions of the commercial bed by the ratio of the kinematic viscosities of the gas raised to the two-thirds power. By taking the ratio of Reynolds number based on the particle diameter to Reynolds number based on the bed diameter

Eq. (70)

ρ f uo L µf     ⋅ =  L  =  L  µf ρ f uo d p  d p m  d p c

The particle diameters in the model scale by the same factor as the bed diameter, by the ratio of the kinematic viscosities to the two-thirds power. Equating the Froude number and rearranging, 1

Eq. (71)

uom  2 =  Lm  uoc  Lc 

Thus, the velocity scales are the square root of the linear dimension scale. By satisfying both Eq. (69) and Eq. (71), the Reynolds number and the Froude numbers are kept identical between the model and the commercial bed. Combining Gs /ρs uo and the product of Reynolds and Froude number along with Eq. (69), it can be shown that

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Eq. (72)

 Gs     ρs m =  ν f m νf  Gs   c    ρs c

1

3   

so that the ratio of solids-flow to solids-density scales as the ratio of the cube root of the kinematic viscosity. Once the model fluid and its pressure and temperature are chosen, which sets the gas density and viscosity, there is only one unique set of parameters for the model which gives similarity when using the full set of dimensionless parameters. The dependent variables, as nondimensionalized by Eq. (18), will be the same in the respective dimensionless time and spatial coordinates of the model as the commercial bed. The spatial variables are nondimensionalized by the bed diameter so that the dimensional and spatial coordinates of the model is proportional to the twothirds power of the kinematic viscosity, as given by Eq. (69)

Eq. (73)

xm xc

 ν fm =   ν fc

2

3   13

Since the velocity scales with vf , the ratio of time scales can be expressed as 1

Eq. (74)

tm = ν fm 3  νfc  tc  

Similarly, it can be shown that the frequency scales as 1

Eq. (75)

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fM fc

 ν fc  3  =  ν fm 


59

Table 1 gives the values of design and operating parameters of a scale model fluidized with air at ambient conditions which simulates the dynamics of an atmospheric fluidized bed combustor operating at 850oC. Fortunately, the linear dimensions of the model are much smaller, roughly one quarter those of the combustor. The particle density in the model must be much higher than the particle density in the combustor to maintain a constant value of the gas-to-solid density ratio. Note that the superficial velocity of the model differs from that of the combustor along with the spatial and temporal variables. When modelling a pressurized hot bed (Table 2) the ambient temperature model fluidized with air has dimensions very close to those of the pressurized combustor. If another gas is used in the model, particularly a gas with a higher density, the model can be made much smaller than the pressurized combustor (see Table 3). Care must be taken to select a safe modelling gas and one which yields a solid density for the model which is available.

Table 1. Atmospheric Combustor Modelled by a Bed Fluidized with Air at Ambient Conditions

GIVEN: Temperature (°C)

Commercial Bed 850

Scale Model, full scaling laws 25

4.45

1.81

Density (kg/m ) Derived from Scaling Laws: Solid Density

0.314

1.20

ρ sc

3.82 ρ sc

Bed Diameter, Length

Dc

0.225D c

Particle Diameter

d pc

0.225d pc

Superficial Velocity

u oc

0.47u oc

(G s /r s ) c

0.47(G s /r s )c

tc fc

0.47t c 2.13f c

-5

Gas Viscosity (10 kg/ms) 3

Volumetric Solid Flux Time Frequency

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60


Table 2. Pressurized Combustor Modelled by a Bed Fluidized with Air at Ambient Conditions

GIVEN: Temperature (°C) -5


Density (kg/m ) Pressure (bar) Derived from Scaling Laws: Solid Density Bed Diameter, length Particle Diameter Superficial Velocity Volumetric Solid Flux Time Frequency

Commercial Bed 850


4.45

1.81

3.14 10

1.20 1

ρsc Dc d pc u oc (Gs /ρs)c tc fc

0.382 ρsc 1.05D c 1.05 d pc 1.01 u oc 1.01(Gs/ρs)c 1.01t c 0.98f c

Table 3. Pressurized Combustor Modelled by a Bed Fluidized with Refrigerant Vapor 134a at Ambient Conditions

GIVEN: Temperature (°C) -5


Density (kg/m ) Pressure (bar) Derived from Scaling Laws: Solid Density Bed Diameter, length Particle Diameter Superficial Velocity Volumetric Solid Flux Time Frequency

9-Oct-2001 JMR

Commercial Bed 850


4.45

1.19

3.14 10

4.34 1

ρsc Dc d pc u oc (Gs/ρs )c tc fc

1.38 ρsc 0.334D c 0.334 d pc 0.58 u oc 0.58(Gs /ρs)c 0.58t c 1.7f c


61

Design of Scale Models Using the Simplified Set of Scaling Relationships

The simplified scaling relationships, Eq. (53), offer some flexibility in the model design since fewer parameters must be matched than with the full set of scaling relationships. When the fluidizing gas, the pressure and temperature of the scale model are chosen, the gas density and viscosity for the scale model are set. The model must still be geometrically similar to the commercial bed. There is still one free parameter. Generally this will be the linear scale of the model. For the simplified scaling relationships, the gas-to-solid density ratio must be maintained constant

Eq. (76)

 ρf   ρf    =   ρs   ρ s   m  c

With ρf of the model set by the fluidized gas and its state, the solid density in the model follows from Eq. (76). Choosing the length coordinate of the model, Lm , which is now a free parameter, the superficial velocity in the model is determined so that the Froude number remains the same,

Eq. (77)

 u 2o    =  gL m

 u 2o     gL c

so that 1/2

Eq. (78)

u om =  Lm    uoc  Lc 

Note that in simplified case, the velocity scaling is not uniquely tied to just the gas properties as it is in the full scaling relationship. With uo and ρs set, the solids recycle rate can be determined by

Eq. (79)

   Gs    =  Gs  ρ ρ u u o o  s m  s c

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62


1

Eq. (80)

G sm =  ρsm   uom  =  ρ fm   Lm  2      Gsc  ρsc   uoc   ρ fc   Lc 

Finally the mean particle size for the model as well as the sphericity and particle size distribution must be determined. The particle size is determined by the need for equal values of uo /umf between the model and the commercial bed.

Eq. (81)

 uo   umf

Eq. (82)

(umf )m = (umf )c  uuom  = (umf )c  Lm   oc   Lc 

  uo  =  m  umf

  c 1/2

In general, umf is a function of the particle diameter and gas properties, as well as φ and εmf . Once the fluidizing gas and the length of scale of the model is chosen, the proper particle diameter is that which gives the value of umf needed in Eq. (82). If both the model and commercial bed are in the region where the respective Reynolds numbers based on particle diameter and gas density are very low, then a single algebraic relationship can be developed. In that region

ρs d 2p µ

Eq. (83)

u mf ~

Eq. (84)

umfm  d pm  ρsm µ c  2  =  = uom =  Lm  umfc uoc  Lc   d pc  ρsc µ m

2

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1


63

1

2

Eq. (85)

2  d pm  ρ µ   = sc m  Lm   d pc  ρsm µc  Lc 

Eq. (86)

2 ρ f µ  L 1/2  d pm    = c m  m  d pc   ρ f m µ c  Lc 

when both (Redp)m and (Redp)c < 20. When the Reynolds number of the model and commercial bed are both very large Eq. (87)

2 u mf ~ d p

ρs ρf

Since the gas-to-solid density ratio of the model and the commercial beds must be the same to satisfy the simplified scaling relationships, Eq. (87) combined with Eqs. (81) and (78) becomes,

Eq. (88)

u 2mfm 2 umf c

=

dpm d pc

 ρs   ρf   ρs   ρf 

   2 m = uom = Lm 2 Lc  uoc   c

When both (Redp)m and (Redp)c > 1000

Eq. (89)

d pm Lm = d pc Lc

Tables 4 and 5 show the values of the mean particle diameter for models of an atmospheric and pressurized commercial bed, respectively, for different selected linear scale ratios between the model and commercial bed. By the use of the simplified scaling parameters, the linear scale factor can be changed as shown in Tables 4 and 5. Note that as the linear

9-Oct-2001 JMR

64


scale factors are changed the particle diameters change much more slowly. The scale model of the 10-atmospheres bed has a mean particle diameter which is quite close to the mean particle diameter of the commercial bed. The model particles have a substantially lower density in this case. It is not clear where cohesive forces will become important. The use of very dense particles (for the models of the one atmospheric bed) will cause a shift of the boundary of cohesive influence as given, for example, by Geldart’s classification. However, adequate experimental data is still lacking with such dense fine particles to definitely set the limits of cohesive influence. Note that for completeness, the nondimensional particle size distribution, sphericity and the internal angle of friction (for slugging and spouting beds) should also be matched between the two beds.

Table 4. Scale Models of Atmospheric Commercial Hot Bed Using the Simplified Scaling Relationship

Commercial Beds

Particle Diameter of Model with Bed Linear Scale Factor

dp

umf

Lm/Lc = 1/4

Lm /Lc = 1/9

40 µm

7.45 × 10 -4 m/s

10 µm

8 µm

60

1.68 × 10-3

15

12

100

4.66 × 10

-3

24

20

200

1.86 × 10-2

49

40

400

7.42 × 10-2

98

80

1000

0.441

245

198

Commercial Bed: T = 800oC, P = 1 atm, ρs = 2500 kg/m3, gas:air Model Bed: ρs = 8960 kg/m3, gas:air at STP

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65

Table 5. Scale Model of 10 Atm Commercial Hot Bed Using the Simplified Scaling Relationships

Commercial Beds

Particle Diameter of Model with Bed Linear Scale Factor

dp

umf

Lm/Lc = 1/4

Lm/Lc = 1/9

40 mm

7.44 × 10-4 m/s

31 mm

25 mm

60

1.67 × 10-3

46

38

100

4.65 × 10-3

78

63

200

1.85 × 10-2

155

126

400

7.18 × 10-2

310

250

1000

0.329

610

550

Commercial Bed: T = 800oC, P = 10 atm, ρs = 2500 kg/m3, gas:air Model Bed: ρs = 896 kg/m3, gas:air at STP

8.0

EXPERIMENTAL VERIFICATION OF SCALING LAWS FOR BUBBLING BEDS

8.1

Hydrodynamic Scaling of Bubbling Beds

Experiments using scaled models of bubbling beds have been carried out since 1980 using the scaling relationships presented in previous sections. The earliest work was concerned with experimental testing and verification of the scaling relationships. Hot bed behavior was compared to cold scaled models in several studies. In others, cold beds of different sizes were compared to each other. For bubbling beds, verification of the scaling relationships needs to include direct or indirect measurements of bubble characteristics. Since bubbles are the prime motive agents for both gas and solids displacement, a detailed verification should involve comparison of bubble properties throughout the large bed and the

9-Oct-2001 JMR

66


scale model. This is preferable to comparison of overall performance parameters for a bed based solely on input and exhaust measurements. For beds fluidized with air at ambient conditions, a number of direct techniques to measure or infer bubble properties exist including local capacitance probes, light reflection, light transmission, videos of bed walls and even computer tomographic techniques. These local methods are difficult, if not impossible to use in a high temperature reactor or combustor. Alternatively, many investigators have used pressure fluctuations to infer bubble characteristics. Lirag and Littman (1971) used statistical analysis of pressure fluctuations to estimate the average size of bubbles leaving a bubbling bed. Fan et al. (1981) concluded that bubble motion and coalescence are responsible for pressure fluctuations in fluidized beds. They also found that the amplitude of the pressure fluctuations is related to bubble size. Therefore, pressure fluctuation measurements, when properly made, have been shown to reflect bed hydrodynamics. For local bed behavior, the differential pressure measurements should be made over a modest bed level rather than using a single pressure tap in the bed with the other tap in the freeboard. Differential pressure measurements reflect local conditions. Roy and Davidson (1989) found that the maximum pressure difference between two closely spaced pressure taps is related to the bubble diameter at that level in the bed. Use of a single bed pressure point can be difficult to interpret since pressure fluctuations can be due to local effects, bubbles erupting at the bed surface, or even fluctuations in the supply system. Roy and Davidson found that pressure fluctuations measured using a single pressure point are primarily due to surface bubbles. They also concluded that the dominant frequency and amplitude measured at a single pressure point are independent of position in the bed. Due to the existence of pressure waves from many locations in a gas fluidized bed, Bi, Grace and Zhu (1995) concluded that absolute probes are not suitable for determining local bubble behavior. Although complete bubble characteristics cannot be determined from differential pressure fluctuations, if the time-resolved differential pressure signal from several locations in a large bed are identical to measurements at corresponding location in the scale model, it is likely that the local characteristics of bubbles are similar. Table 6 presents the range of scaling parameters for tests undertaken by different investigators to verify the scaling relationships for a bubbling bed.

9-Oct-2001 JMR

Table 6. Scaling Parameter Values for Bubbling Fluidized Bed Experimental Studies

Fluidized Bed Scale-up 67

9-Oct-2001 JMR

~~~

..,~ ~~

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~~ ..;..;

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--r-o..r--

::;;~ ~N

~~~

~~ -0-0

~~

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~~

ii

~~~ ;"

~~i

}kk

ii

k}

~~1~~

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::!=1$~

~~ ,-;,-;

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~~

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..*' -ON .,.;.,.;

1

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;;;g~~~ ..,~..,~.., ~@§~~

~~~ ;;;:5:~ i ..,..,-

"0"0"0"0"0 RRRa~

]

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~!f i 1~ o~ 15

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.-:,.,

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-:""!,.,a~ .."".'-1-",

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r-~~~-o ~~"'~~~

~~§~~~ ooN.';oo

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68

9-Oct-2001 JMR

Table 6 (cont’d.)


69

Verification of Scaling Relationships for Bubbling and Slugging Beds

Most early experiments devoted to verifying the scaling relationships have dealt with the full set of scaling relationships. Several more recent experiments have dealt with a reduced set of dimensionless parameters. In some experiments, additional scaling parameters were unintentionally matched. Fitzgerald and Crane (1980) were one of the first to evaluate the full set of hydrodynamic scaling parameters. They compared the hydrodynamics of two scaled beds using pressure fluctuation measurements and movies. In one bed, cork particles were fluidized with air; the other bed used sand fluidized with pressurized refrigerant 12 vapor. Movies showed qualitative agreement between bubble growth and the solids flow in the beds. The ratio of the bed minimum fluidization velocities was within 20% of the theoretical value; the difference was attributed to the angular shape of the cork particles, φS was not matched between the beds. The fast Fourier transform of the pressure fluctuations was used to determine the average frequency of the fluctuations which should be related to the bubble frequencies. The ratio of average frequencies for the two beds was in fair agreement with the theoretical velocity-time scale factor. The pressure fluctuation data were taken with a single bed pressure tap. This may be responsible for the level of agreement in the frequency ratios. Some additional qualitative slugging comparisons were made using movies of bed behavior; the slugs appeared to have the same scaled lengths and velocities. Fitzgerald et al. (1984) measured pressure fluctuations in an atmospheric fluidized bed combustor and a quarter-scale cold model. The full set of scaling parameters was matched between the beds. The autocorrelation function of the pressure fluctuations was similar for the two beds but not within the 95% confidence levels they had anticipated. The amplitude of the autocorrelation function for the hot combustor was significantly lower than that for the cold model. Also, the experimentally determined time-scaling factor differed from the theoretical value by 24%. They suggested that the differences could be due to electrostatic effects. Particle sphericity and size distribution were not discussed; failure to match these could also have influenced the hydrodynamic similarity of the two beds. Bed pressure fluctuations were measured using a single pressure point which, as discussed previously, may not accurately represent the local hydrodynamics within the bed. Similar results were

24-Sep-2001 JMR

70


obtained between two two-dimensional beds: a bed of reacted limestone fluidized with helium and a half-scale bed of copper fluidized with air. Nicastro and Glicksman (1984) experimentally verified the full set of scaling laws for bubbling fluidized beds. They compared the timeresolved differential pressure measurements from a bubbling fluidized bed combustor and a scaled cold model. Good agreement was obtained between the spectral content and the probability density distribution of the differential pressure fluctuations, indicative of the bubble size distribution, between the hot combustor and the cold model. Figure 29 presents the comparisons. They concluded that hydrodynamic similarity had been achieved between the hot combustor and the cold model. The solid-to-gas density ratio, which was not matched exactly in the comparison, differed by 23%, but the Reynolds number based on particle diameter (Redp) was approximately 5, suggesting the beds were operating in the viscous limit. At low particle Reynolds numbers, viscous forces are more important than fluid inertial forces making the density ratio a less important scaling parameter. When actual hot bed material was used in the cold model, a violation of the scaling laws, the model’s behavior was very different from that of the hot bed. Horio et al. (1986a) used three geometrically similar bubbling beds, fluidized with ambient air, to verify their proposed scaling laws. This has been shown by Glicksman (1988) to be equivalent to the viscous limit scaling laws, Eq. (66). The solid-to-gas density ratio was not varied in the experiments although it was not one of the proposed scaling parameters. By maintaining a constant density ratio, they, in essence, used the simplified set of scaling parameters, Eq. (53). Video analysis of bubble eruptions at the bed surface were used to determine the crosssectional average bubble diameter, bubble diameter distribution, and radial distribution of superficial bubble velocity. Similarity was achieved in these hydrodynamic parameters when bed Froude number, density ratio, and the ratio of superficial to minimum fluidization velocities were matched. Horio et al. (1986b) verified the bubbling bed scaling laws of Horio et al. (1986a) for solid mixing and segregation. Sand was used as a bed material in straight and tapered bed geometries. A bed sectioning technique was used to measure the transient radial dispersion coefficient and the distribution of float tracers. They concluded that bed mixing and the behavior of floating bodies obey the scaling laws in both straight and tapered beds. The solid-to-gas density ratio was again held constant in the tests, satisfying the simplified set of scaling laws. The particle Reynolds numbers (Redp) were approximately unity or less.

24-Sep-2001 JMR


0

2

3

4

71

x 10-3

Frequency, fl (uoldp )

Figure 29. Comparison of dimensionless power spectra of differential pressure fluctuations. Double probe across levels 2 and 3; x/L = 0.0, coal burning bubbling bed combustor. Full set of scaling laws with iron grit in cold bed; hot bed material in cold bed violates scaling laws. (From Nicastro and Glicksman, 1984.)

Newby and Keairns (1986) made bubbling bed scaling comparisons between two cold models using the full set of scaling laws. One bed was fluidized with two different 200 µm glass powders using ambient air. The second bed, which was a half-scale model of the first, used pressurized air to fluidize 100 µm steel powder. High-speed movies showed good agreement between the nondimensional bubble frequencies in the two beds. Figure 30 is a plot of the nondimensional bubble frequencies as a function of bed Froude number. They also found reasonably good agreement between the nondimensional amplitudes of the pressure fluctuations in the beds.

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~ ~ m k.

Figure 30. Comparison of non-dimensional bubble frequencies from two cold scaled models. (Newby and Keairns, 1986.)

Zhang and Yang (1987) carried out scaling comparisons between two two-dimensional beds with uo2/gD and uo /umf matched between them. They also inadvertently kept the solid-to-gas density ratio constant; thus they matched the simplified scaling parameters. They found through photographs that the beds appeared qualitatively similar. The beds also had similar dimensionless freeboard entrainment rates and dimensionless bed heights over a range of uo /umf . Roy and Davidson (1989) considered the validity of the full and viscous limit scaling laws at elevated pressures and temperatures. The nondimensional dominant frequency and amplitude of the pressure drop fluctuations were used as the basis of the comparison. They concluded that when the full set of scaling parameters is matched, similarity is achieved. They also suggested that it is not necessary to match the density ratio (ρs /ρf) and dp /D, the simplification for viscous limit scaling, for particle Reynolds numbers (Redp) less than 30. Although the only run with Redp near 30 which was similar to the low Reynolds number test had the same density ratio as the low Redp runs. These conclusions may be open to different interpretations. As shown in Table 6, the scaling parameters neither matched closely nor varied in a systematic manner.

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Di Felice et al. (1992a) investigated the validity of the full set of scaling laws for bubbling and slugging fluidized beds. They used an experimental facility which permitted the pressurization of different diameter test sections to match the scaling parameters. Minimum fluidization measurements, video measurements of bed expansion, and pressure fluctuation data were used to compare the similarity of five different bed configurations. Three of the beds were scaled properly, the fourth had a mismatched particle sphericity, and the fifth bed was purposefully misscaled relative to the others (see Table 6). The voidage at minimum fluidization was found to be the same for all the beds except the one with the different particle sphericity. In the bubbling regime, good agreement in the nondimensional bed expansion measurements was obtained for all but the bed with the mis-scaled particle sphericity. The lower particle sphericity increased umf for the system which effectively shifted the bed expansion curve for this case. The pressure fluctuations for the three properly scaled beds in the bubbling regime showed good agreement while the mis-scaled beds exhibited poor agreement with the other three. Figure 31 is a plot of the dimensionless dynamic pressure variance for the five beds in the bubbling regime. The two sets of data which deviate from the other three correspond to the mis-scaled beds.

0.03

b. 0.02

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.

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.

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~: M

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"0'"t Figure 31. Comparison of dynamic pressure variance for three properly scaled beds and two mis-scaled beds in bubbling regime (DiFelice, et al., 1992a). Properly scaled: ¨, laposorb; ∆ , sand; ¡, bronze. Intentionally mis-scaled: +, iron; \, sand.

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In the slugging regime, Di Felice et al. (1992a) found that the bed expansion characteristics were similar to those in the bubbling regime, but the pressure fluctuation characteristics for all five beds were in poor agreement with each other. Figure 32 is the plot of the dimensionless dynamic pressure variance for the five beds in the slugging regime. They attributed this to the importance of particle material properties and some particle-particle interaction effects which are not accounted for in the full set of scaling laws. This is discussed further below in the light of recent results for spouting beds.

*... ~

uo'ut Figure 32. Comparison of dynamic pressure variance for three properly scaled beds and two mis-scaled beds in slugging regime (DiFelice et al., 1992a). Properly scaled: ¨, laposorb; ∆ , sand; ¡, bronze. Intentionally mis-scaled: +, iron; \, sand.

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Di Felice et al. (1992b) evaluated the full set of scaling laws for three different Geldart powder categories (A, B, and D) in the bubbling and slugging fluidization regimes. Pressure fluctuations were used as the basis for the scaling comparisons. In the bubbling regime, the RMS and dominant frequencies of the pressure fluctuations showed good agreement for all three powder categories. Only Geldart groups B and D were considered in the slugging regime. They exhibited fair agreement in the RMS of their pressure fluctuations, but their dominant frequencies disagreed. They found that the full set of scaling laws are valid for bubbling beds fluidizing powders in Geldart groups A, B, and D. They also concluded that the full set of scaling laws is not sufficient for slugging beds where particle-particle interactions are also thought to be important. 8.3

Verification of Scaling Laws for Spouting Beds

He et al. extended the scaling considerations to spouting beds. They showed that for spouting beds the full set of scaling relationships, Eq. 37, must be augmented with two new parameters, the internal friction angle and the loose packed voidage. By systematic tests in different sized cold beds as well as comparisons between hot and cold beds, they showed excellent agreement when the full set of scaling parameters augmented with the two spouting bed parameters were held constant. Close agreement was found for spout diameter, fountain height, longitudinal pressure profiles and dead zone boundary. Figure 33 shows a comparison of dimensionless spout diameter versus dimensionless height. Cases A and B differ in bed diameter by a factor of 2 but all of the dimensionless groups are maintained equal. In case C, the Reynolds numbers based on particle diameter and bed diameter are mismatched and in case D, the Froude number is mismatched. When the internal friction angle and sphericity were mismatched, there was a large disagreement in fountain height. The internal friction angle is also important for slugging beds (Zenz and Othmer, 1960). DiFelice et al. (1992 a, b) did not report their values; it could be that the disagreement they found in their slugging bed tests was due to mismatches of the internal friction angle.

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J: ':8, .c 0) .0) .c f/) f/) ~ c o .0 c 0) .-E 0

Figure 33. Dimensionless spout diameters as a function of dimensionless height for small columns. Case A: test case; Case B: all dimensionless parameters matched, bed diameter halved; Case C: particle Reynolds number mismatched; Case D: Froude number mismatched; Case E: density ratio, Reynolds number mismatched; Case F: bed Reynolds number mismatched; Case G: internal friction angle, loose packed voidage mismatched. (From He et al., 1995.)

8.4

Verification of Scaling Relationships for Pressurized Bubbling Beds

Almstedt and Zakkay (1990) made scaling comparisons between a hot pressurized fluidized bed combustor with horizontal tubes and a pressurized cold scale model using the full set of scaling laws. The cold model had linear dimensions one-half those of the hot bed. A capacitance probe was used to measure the mean values of the bubble frequency, pierced length, bubble rise velocity, and bubble volume fraction. Scaling comparisons were made using the dimensionless form of these dependent

9-Oct-2001 JMR


77

hydrodynamic parameters. Three different bed materials were used in the cold bed: Olivine sand and two different size distributions of the hot-bed material, one properly scaled and one out of scale. The solid-to-gas density ratio of the sand was 17% higher than the ratio for the combustor while the scaled down hot bed material, when used in the cold bed, has a solid-to-gas density 14% lower than the ratio for the combustor. The sand also had a lower sphericity than the hot bed material. The out-of-scale hotbed material was also used to illustrate the sensitivity of the scaling to the D/d p parameter. The nondimensional form of the capacitance probe measurements agreed within 25% for the sand and the properly scaled hotbed material; the agreement was best in the upper part of the bed. The dimensionless bubble velocity disagreed by as much as 18%. The properly scaled hot-bed material showed only slightly better agreement than that for the sand, but the mismatch in the density ratio and the sphericity for the sand was small. The improperly scaled hot-bed material had a maximum deviation of 38% from the hydrodynamics of the hot-bed combustor. When the pressure or the superficial velocity of the cold bed was changed from the correct value based on the scaling parameters, the disagreement in bubble properties increased. Almstedt and Zakkay concluded that behavior which is hydrodynamically similar to that of a pressurized fluidized bed combustor can be achieved using a properly scaled cold model. Glicksman and Farrell (1995) constructed a scale model of the Tidd 70 MWe pressurized fluidized bed combustor. The scale model was fluidized with air at atmospheric pressure and temperature. They used the simplified set of scaling relationships to construct a one-quarter length scale model of a section of the Tidd combustor shown in Fig. 34. Based on the results of Glicksman and McAndrews (1985), the bubble characteristics within a bank of horizontal tubes should be independent of wall effects at locations at least three to five bubble diameters away from the wall. Low density polyurethane beads were used to obtain a close fit with the solid-to-gas density ratio for the combustor as well as the particle sphericity and particle size distribution (Table 6). Differential pressure measurements were made between several vertical elevations within the bed. The probability density function of the cold model and combustor gave very close agreement (Fig. 35). The solid fraction profiles were obtained from the vertical pressure profile with a hydrostatic assumption. The cold model solid fraction profile showed very close agreement with data taken from pressure taps in two different locations within the combustor (Fig. 36). The solid fraction shows a

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somewhat unexpected behavior with a lower value near the distributor, possibly due to many small low velocity bubbles at that location which coalesce and increase in velocity until the tube bank is encountered. The power spectral density of the combustor exhibited several distinct peaks at increasingly higher frequencies. All but the first peak were not seen in the cold model. The peaks could be due to tube vibrations in the hot bed, fluctuations upstream or downstream of the bed or hydromechanical interactions between the bed and the internals. The long length of the pressure leads or the in-bed location of the taps could also have contributed to the peaks. Further measurements in the combustor are needed to resolve their origin.

Figure 34. Tidd boiler enclosure with scaled section designated. (From Glicksman and Farrell, 1995.)

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Figure 35. Comparison of the Tidd PFBC and cold model based on simplified scaling laws at y/H = 0.34. (From Glicksman and Farrell, 1995.)

BedS,

.8 ',-

i 1 ~ .4

-, ,

m ""',. --',

9 G G.--0

', , 'c

V ~ ~ TE] ~ eJ!ll. T~~

',,:~ ...

v ,

I I

'.

I ...I ..',1

.2

,.'

~~ c'

,~'m ..R

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.3

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Figure 36. Solid fraction profile comparisons for Tidd PFBC and cold model based on simplified scaling laws. (From Glicksman and Farrell, 1995.)

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Farrell (1996) experimentally evaluated the importance of the solid-to-gas density ratio (ρS /ρf) for scaling the hydrodynamics of bubbling and slugging fluidized beds. Two bed materials, polyethylene plastic (ρs = 918 kg/m 3) and a dolomite/limestone sorbent mixture (ρs = 2670 kg/m3), were used to create a mismatch in the density ratio. The size of the particles was chosen such that the remaining simplified scaling parameters were matched. Measurements showed that the internal angle of friction was similar between the two materials. A 10.2 cm diameter bed with a mock tube bundle was used to make bubbling regime comparisons, while a 5.1 cm diameter bed was used to compare the behavior of the two materials in the slugging regime. In addition, the character of the bubbling-slugging transition was also explored. Figure 37 compares the solid fraction profiles in the 10.2 cm bubbling bed for the two bed materials at uo /umf = 1.2. In this particular case, the solid fraction of the sorbent material is less than the plastic in the lower regions of the bed, with good agreement in the upper section of the bed. However, for the same conditions, the dimensionless standard deviation of the time-varying pressure drop showed the best agreement in the bottom of the bed with a large discrepancy in the upper portion of the bed. The bubbling-slugging transition behavior was evaluated in the 5.1 cm diameter bed. Horio et al. (1992) proposed using the inflection points in plots of the root-mean-square of the gauge pressure versus superficial velocity (uo) to identify flow regime boundaries. Figure 38 shows this behavior for the two bed materials. The plastic bed material has a much broader transition region between its fully bubbling and fully slugging regimes than the sorbent material and the nature of this transition is different between the two materials. Therefore the solid-to-gas density ratio influences both the hydrodynamics in the bubbling regime and the boundary at which the transition to slugging occurs. This is consistent with the conclusion of Glicksman et al. (1993b) where they found that it is essential to match the density ratio when scaling circulating fluidized bed hydrodynamics. 9.0

APPLICATIONS OF SCALING TO COMMERCIAL BUBBLING FLUIDIZED BED UNITS

A substantial number of experimental demonstrations of the validity of scaling has increased awareness of the concept and confidence in its application. Although applications to commercial designs have been undertaken, unfortunately only a modest number have been documented in the open literature.

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Scaling has many useful applications. The dynamic characteristics of different bed designs can be quickly compared. The influence of bed diameter on hydrodynamic behavior can be studied by the use of several different size models. The models allow easy experimental examination of existing operating characteristics. The beds also can be used to quickly confirm the influence of proposed modifications. Since the models usually operate at ambient conditions, it is possible to instrument them to observe detailed behavior. This allows a better understanding of the fundamental physics as well as the identification of hydrodynamic factors needed for proper correlation of performance. The earliest scaling studies were directed at atmospheric bubbling bed combustors. To date, a rich variety of questions have been addressed. Jones and Glicksman (1986) constructed a model of the 20 MW e bubbling bed pilot plant jointly sponsored by the Tennessee Valley Authority and the Electric Power Research Institute (EPRI) at Paducah, Kentucky. Figure 39 shows a photograph of the in-bed tubes installed in the scale model. The model, which is roughly 100 by 120 cm in cross section, simulates twothirds of the entire 20 MW pilot plant. Care was taken to carefully match the pilot plant tube bundle geometry and distributor design. Steel grit particles with the same dimensionless size distribution and sphericity as the hot bed material were used. The full set of scaling parameters was matched in the model and the combustor. The largest discrepancy was in the solid-to-gas density ratio which was 18% smaller in the model than the pilot plant.

Figure 39. Model of 20 MW bubbling fluidized bed combustor showing tube arrangement. (From Jones and Glicksman, 1986.)

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83

Optical probes were used to measure the bubble size, frequency and velocity within the dense bed. The bubble velocity for an actively bubbling bed was found to closely agree with the drift flux form proposed by Davidson and Harrison (1963). In contrast, the volumetric flow rate of the bubbles was found to be far less than that predicted by the two-phase hypothesis (Fig. 40).

Data 1.

Set

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0.8

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(m/s)

Figure 40. Deviation from two-phase theory for model of bubbling fluidized bed combustor.(From Jones and Glicksman, 1986.)

The larger particles were thrown high in the splash zone; higher than predicted by a ballistic trajectory using the bubble rise velocity as the initial velocity and neglecting any air drag. Later observations of this model showed that, when bubbles erupt at the surface, the accompanying

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gas flow has a velocity much higher than the bubble rise velocity (Glicksman and Piper, 1987). This led to a mechanistic model for gas throughflow aided by the low resistance of the bubble cavity (Yule and Glicksman, 1988) and an accurate prediction of bubble volume flow rate and bed expansion (Glicksman et al. 1991b). Commercial bubbling bed combustors must operate satisfactorily over a range of part-load conditions. While reducing the total combustion rate, it is necessary to keep bed operating temperature constant. This requires a reduction in the heat transfer to the water-filled tubes within the bed. One technique utilizes the contraction of the bed which accompanies a decrease in superficial velocity. As the bed contracts, some of the tube rows are uncovered, reducing the net heat transfer. The scale model allowed many different tube arrangements to be tested; six tube configurations were examined. Figure 41 shows three of the six different tube bank configurations which were tested. The validity of the scaling technique was confirmed by a comparison of the bed expansion measured for the pilot plant and that found in the model equipped with the same tube bank geometry (Fig. 42). A second method to reduce load while maintaining constant bed temperature is to reduce the superficial velocity below umf to a portion of the bed. In this design, the bed does not contain vertical partitions above the distributor. The scale model was used to determine the rate of growth of the fixed bed in the defluidized zone along with the heat transfer to tubes in that region. Figure 43 shows a typical pattern of particle accumulation in a slumped zone adjacent to an actively fluidized zone. Heat transfer coefficients are also shown. Note that tubes near the upper surface of the defluidized zone, which experience a downflow of solids, have a very high heat transfer rate. In a bubbling bed operating at high ratio of uo /umf , there is a considerable amount of solids present in the freeboard, particularly near the bed surface in the so-called “splash zone.” The high density of particles in the freeboard can cause substantial combustion and emission release in that zone as well as freeboard overheating if tubes are not present. The average density of solids was measured in the freeboard of the scale model of the 20 MW pilot plant (Glicksman and Piper, 1987). As shown in Fig. 44, the average density in the freeboard decreases exponentially with distance above the dense bed. Also shown on the figure is the predicted behavior based on a bubble eruption model (Glicksman and Yule, 1991).

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Figure 41. Tube configuration used for bed expansion studies. Note: all dimensions in cm scaled to equivalent sizes in the combustor. (From Glicksman et al., 1991b.)

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Figure 42. Bed expansion for TVA 20 MWe FBC and for scale model, shown in Fig. 39. (From Glicksman et al., 1989.)

Figure 43. Particle accumulation in slumped zone adjacent to active bed: air velocity through active region = 151 cm/s; air velocity through slumped region = 7.3 cm/s; numbers above heater are heat transfer coefficient in W/m2 °C after 15–30 min. (From Glicksman et al., 1992.)

9-Oct-2001 JMR


O

3

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Height Y/Db Figure 44. Density in freeboard of bubbling bed, present theory with Db equal to horizontal tube pitch, 3.9 cm. (Data of Glicksman and Piper, 1987.)

The heat transfer from tubes in the freeboard was also measured for the 20 MW model. Figure 45 shows a comparison of the measured overall heat transfer coefficient in the 20 MW pilot plant versus that predicted from the scale model test. When the bed height is lowered, uncovering some tubes, the heat transfer is reduced because there are fewer particles contacting the tube surface. Although the scale model did not include proper scaling for convective heat transfer, the rate of change of the overall heat transfer should be a function of the hydrodynamics. Ackeskog et al. (1993) made the first heat transfer measurements in a scale model of a pressurized bubbling bed combustor. These results shed light on the influence of particle size, density and pressure levels on the fundamental mechanism of heat transfer, e.g., the increased importance of the gas convective component with increased pressure.

9-Oct-2001 JMR

88


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Tsukada et al. (1991) applied the CFB scaling laws of Horio et al. (1989) at several different elevated pressures (viscous limit scaling laws). A single bed and bed material were used in the study. A pressure vessel was used to vary the gas pressure. The bed was fluidized with ambient temperature air at three different pressures (0.1, 0.18, and 0.35 MPa). Axial solid fraction profiles and optical probe measurements were used as the basis for their similarity comparison. They found that as the pressure was increased, the axial solid fraction profile changed, indicating a change in the hydrodynamics. It was suggested that the effect on the axial solid fraction profile could be due to reaching a Reynolds number limit, e.g., the upper boundary for the viscous limit. They also suggest that it could be due to a change in gas bypassing between the riser and the downcomer. In this study, the gas-to-solid density ratio changes with pressure level. It is likely based on the recent results of Glicksman et al. (1993b) that the change in gas-to-solid density ratio led to the influence of the pressure level on the bed hydrodynamics. This points out the deficiencies of the viscous limit, the Horio et al. (1989) scaling relationship.

9-Oct-2001 JMR


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Glicksman et al. (1991) made scaling comparisons between an experimental circulating fluidized bed combustor and a scaled cold model based on the full set of scaling laws. The time-resolved pressure fluctuations and the time-averaged pressure drop were measured. Due to uncertainties in the hot bed solid circulation measurements, the cold bed solids flux was adjusted until the average bed solid fraction matched that of the hot bed. Differences could also be due to the mismatch in the solid-to-gas density ratio between the model and the hot bed. The vertical solid fraction profiles, the probability density function and the Fourier transform of the pressure fluctuations were compared between the hot and cold bed. Good agreement was obtained between the vertical solid fraction profiles except near the top of the beds. It was suggested that the differences in the solid fraction profiles at the top of the bed could be due to protrusions or wall roughness in the hot bed which were not modeled in the cold bed. Good agreement was also obtained in the comparison of the probability density distribution and the Fourier transform of the pressure fluctuations. Chang and Louge (1992) carried out tests on a circulating bed in which they could vary the gas composition. By combining this with particles of different density and size, they were able to scale a series of different size hot commercial beds with diameters up to five times larger than the cold bed. Comparisons between glass and plastic particles showed identical mean vertical solids fraction profiles. The corresponding pressure fluctuations for plastic and glass are found to scale with ρs gφ dp ; one would expect the pressure fluctuations to scale with ρsuo2. This is probably an artifice of the experimental design since Froude number based on particle diameter, uo2/gφ dp , was matched in the comparisons whereas Froude number based on bed diameter could not be matched since the experimental bed diameter was fixed in the tests. The inability to alter the bed diameter also made it impossible to match the D/ dp scaling parameter. Chang and Louge matched a modified form of the full set of scaling laws. Particle sphericity is not explicitly included as an independent parameter, rather it is included with the particle diameter based on a combination of the gas to particle drag coefficient. Their modified parameters are: Eq. (90)

Fr* =

2 uo gφ d p

, L* =

D dp φ

,M =

Gs ρs u o

,R =

ρs ρf

, and Ar * =

( )3 g

ρs ρ g dp φ

µ2

9-Oct-2001 JMR

96


The values of the parameters matched for scaling are presented in the “Other Parameters” column of Table 7. Steel and glass particles were also compared. The similarity using steel and glass was poor because the bed using the steel particles was choked while the bed with glass particles was not. Yang’s (1983) correlation indicates that choking is a strong function of the Froude number based on bed diameter (FrD). The value of FrD could not be matched between the beds which caused them to choke under different conditions. Glicksman et al. (1993a) evaluated the full set of scaling laws for circulating fluidized beds. Solid fraction data were obtained from the 2.5 MWth Studsvik atmospheric CFB prototype. The full set of scaling laws were evaluated through solid fraction profile comparisons between Studsvik and a 1/4 scale cold model. Fairly good agreement was obtained; the profiles most closely matched in the top of the beds. Differences between the profiles were attributed to uncertainty in the hot bed solid flux measurements and to the mismatch in the solid-to-gas density ratio. The viscous limit scaling laws was also evaluated by Glicksman et al. (1993b) in a series of comparison tests using circulating beds. To compare two cases based on the viscous limit while deliberately mismatching the solid-to gas density ratio, scaling was attempted between glass/steel and glass/plastic (i.e., different density ratios) in the same bed. The average solid fraction profiles, solid fraction probability density functions, and power spectral densities were all in poor agreement. Figure 49 is a sample solid fraction profile comparison based on viscous limit scaling. It is believed the beds were operating near the point of incipient choking condition as predicted by the Yang (1983) correlation. Because this correlation indicates that choking is a strong function of the solid-togas density ratio, the viscous limit scaling parameters are unable to model bed hydrodynamics near the boundary between different flow regimes. Figure 50 shows a somewhat better agreement near the top of the bed at higher velocities. They concluded that since low uo is required for the viscous limit scaling to be valid while sufficiently high uo is required to prevent choking, the applicability of the viscous limit scaling parameters for circulating fluidized beds is limited. It was suggested that these scaling parameters may have a wider range of validity in bubbling beds. The simplified scaling laws were used by Glicksman et al. (1993b) to compare two geometrically similar beds, one having linear dimensions four times larger than the other. In one series of tests, properly sized plastic particles were used in both beds; in another test series, glass particles were used in the two beds. The average solid fraction profiles showed excellent agreement (Figs. 51 and 52). The probability density functions and power

9-Oct-2001 JMR


97

spectral densities also agreed well. In contrast to the viscous limit scaling results, the simplified scaling laws which include the gas-to-solid density ratio as a scaling parameter gave good agreement even for conditions where Yang’s (1983) correlation predicted the bed was choked.

Figure 49. Solid fraction profiles, glass/plastic viscous limit scaling; density ratio mismatched, low velocity case. (From Glicksman et al., 1993a.) 100 M = 0.0043 Fr = 1.39 Repl = 16.5 Regl ~ 13.0

~

10

U o ~ 2.5 m/s

c .2 ~ .I: :2 O 1000), Eq. (1) becomes:

Eq. (4)

2 U mf =

(

)

d p ρ p −ρg g 24.5ρ g

Rep, mf > 1000

Because gas viscosity does not vary significantly with pressure, the only parameter in Eq. (3) which changes with pressure is the gas density. However, because ρp for most materials is so much larger than ρg even at elevated pressures, the term (ρp - ρg) essentially does not change with pressure. Therefore, for small particles, Eq. (3) predicts that Umf will not change with pressure, and agrees with experimental findings.

24-Sep-2001 JMR

116 Fluidization, Solids Handling, and Processing In Eq. (4), the only term which is influenced by pressure is ρg, because (ρp - ρg) does not change with pressure. For large particles, the Wen and Yu Equation predicts that U mf should vary as (1/ρg)0.5. Therefore, Umf should decrease with pressure for large particles. This is also what is observed experimentally. The variation of Umf with temperature is also strongly affected by particle size. For large particles, Eq. (4) predicts that U mf will vary as (1/ρg)0.5. Because temperature causes ρg to decrease, Umf should increase with temperature for large particles. For small particles, Eq. (3) shows that Umf should vary as 1/µ. Therefore, Umf should decrease with temperature for small particles, where viscous effects predominate. Calculated Umf versus temperature curves obtained from the Wen and Yu Equation are shown in Fig. 2 for several different particle sizes. As indicated above, the Wen and Yu Equation predicts that Umf increases with increasing temperature for large particles (where turbulent effects dominate), and decreases with increasing temperature for small particles (when viscous effects dominate). 2 Material: Sand Pressure: 0.1 Bar Gauge

Minimum Fluidization Velocity, m/s

3000 microns 1.5

1 2000 microns

0.5 1000 microns

500 microns 0 0

400

800

1200

o

Temperature, C

Figure 2. The effect of temperature on minimum fluidization velocity. (Wen & Yu.)

24-Sep-2001 JMR

Pressure and Temperature Effects

117

The predicted effects of temperature on Umf were experimentally observed by Botterill and Teoman (1980) as shown in Fig. 3. For sand particles 462 microns in diameter, they observed that Umf decreased with temperature. For larger material (2320 microns) in the “transition” region in Fig. 2 (where Umf changes from increasing with temperature to decreasing with temperature), they observed an initial increase in Umf with temperature at low temperatures, which was followed by a decrease in Umf as increasing temperature caused viscous effects to become dominant.

Minimum Fluidization Velocity, mm/s

250

Material: Sand Particle Size: 462 microns

200

150

100

50

0 0

200

400

600

800

1000

800

1000

o

Bed Temperature, C

Minimum Fluidization Velocity, mm/s

1200

1100

Material: Sand Particle Size: 2320 microns

1000

900 0

200

400

600 o

Bed Temperature, C

Figure 3. The effect of temperature on the minimum fluidization velocity. (Botterill and Teoman.)

24-Sep-2001 JMR

118 Fluidization, Solids Handling, and Processing Temperature can also affect fluidized beds if it causes the particles to sinter or become sticky. Siegell (1984) has shown that above a certain critical temperature, the fluidizing gas velocity must be increased above the minimum fluidization velocity in order to maintain fluidization in the bed. Siegell reported that the increase in velocity required to maintain fluidization was linearly proportional to the temperature increase. Increasing bed L/D, and smaller particle sizes, increased the tendency to agglomerate. Compo (1987) also investigated the effect of temperature on the velocity required to prevent defluidization for (i) materials which agglomerated rapidly after their sintering temperature (Ts) was reached (coal, plastics, etc.), and (ii) materials which agglomerated slowly after reaching Ts (calcium chloride, etc.) For the rapidly agglomerating materials, the gas velocity had to be increased rapidly to prevent defluidization after Ts was reached. For the weakly agglomerating materials, a much smaller increase in gas velocity was required to prevent defluidization after Ts was reached. Although the trends of varying temperature and pressure can be predicted by minimum fluidization velocity correlations, sometimes the absolute values of the predictions are significantly in error. Frequently, the primary cause of uncertainty is in determining a representative value for the average particle size to use in the Umf correlation. One technique which can be used to improve upon the accuracy of the prediction of Umf, is to first determine Umf experimentally at ambient conditions. Then, using this experimentally-obtained value for Umf, back-calculate an effective particle size from the correlation. Using this effective particle size, then calculate Umf at actual conditions. This technique can predict Umf more accurately than by using a correlation alone. It substitutes an effective value for the shape factor and the average particle size. However, it does not account for changes in voidage which might occur with changes in temperature and pressure. Yang et al. (1985) proposed another method of determining Umf at elevated temperature and pressure. As with the procedure described above, Umf is first determined experimentally at ambient conditions. This data point is then located on a curve of constant εmf on a plot of

[( Re ) (C ) ] 2 ε

24-Sep-2001 JMR

1/ 3

D ε

mf

[

vs. ( Re )ε /(C D

)ε ]1mf/ 3


119

where Reε and CDε are Reynolds number and drag coefficients, respectively, modified by Barnea and Mizrahi (1973) to apply to multiparticle systems. This technique produced estimates of Umf which agreed extremely well with experimental Umf data taken at pressures up to 6300 kPa. It also predicted the experimental data of Yoshida et al., (1976) and Mii et al., (1973) obtained at temperatures up to 850°C (Fig. 4). Shrivastava et al. (1986) applied this technique to the data of Saxena and Vogel (1977) also with excellent results. 100

50

Yoshida et al. (1976) Catalyst 214 µ m 151 µ m Sand Catalyst 151 µ m Catalyst 151 µ m Mii et al. (1973)

30

mf

e mf = 0.40 e mf = 0.48 e mf = 0.51

10

2

[(Re) e (C D) e]

1/3

Graphite 300 µ m

20

e mf = 0.60 5 3 2

1 0.001

0.002

0.005

0.01

0.02

0.05

[(Re) e /(C D) e]

0.1

0.2

0.5

1

1/3 mf

Figure 4. Comparison of data and Yang et al. procedure for Umf determination.

This success of this procedure implies that εmf does not change significantly with temperature and pressure. However, other researchers (see below) have reported that εmf can change substantially with temperature and pressure for Geldart Group A powders. Therefore, the authors caution that this technique may not be accurate for Geldart Group A particles.

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120 Fluidization, Solids Handling, and Processing Abrahamsen and Geldart (1980) defined Group A powders as those in which U mb /U mf > 1, and Group B powders as those where Umb /Umf = 1. They developed the following equation to predict Umb /Umf.: 2300ρ g µ Exp (0.716F ) U mb = 0.8 0.934 U mf dp g ρ p − ρ g 0.934 0.126

Eq. (5)

0.523

(

)

where F is a weight fraction of 0 to 44 micron particles in the distribution Equation (5) implies that materials may shift from one Geldart Group to another as a consequence of changing system parameters. For example, if a material is in Group B near the Group B/A boundary and either the gas density and/or the gas viscosity is increased, Umb /Umf will increase and the material can “cross” the boundary into Group A. That this can occur has been demonstrated by experiments performed by Varadi and Grace (1978). They found that Umb /Umf increased from a value of 1 to a value much greater than 1 as system pressure was increased over the range of 1 to 22 atmospheres as shown in Fig. 5. Thus, the material moved from Group B to Group A as a consequence of changing system pressure. Equation (5) predicts that Umb /Umf is more sensitive to gas viscosity than gas density. The data of de Vries et al. (1972) agree with this prediction.

1.2

Bed Voidage and Bed Expansion

Voidage in a bubbling fluidized bed is also affected by temperature and pressure. Botterill et al. (1982) reported for Group B materials, that the Wen and Yu Equation generally over-predicts the reduction of Umf with temperature. They found that the overprediction was due to the increase in voidage at minimum fluidization (εmf) with increasing temperature. For Geldart Group D materials, they found that εmf was insensitive to changes in temperature. Pressure has been found to cause the dense-phase voidage (εd) to increase for Group A materials. The effect diminishes with increasing particle size, and for large Geldart Group B and Group D materials, εd is insensitive to pressure. Weimer and Quarderer (1983) in their highpressure experiments, found that over a pressure range of 8 to 69 bar, εd for

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121

66 and I 08-micron activated carbon increasedby 40 and 16%, respectively, as shown in Fig. 6. They reported essentially no effect of pressure on EXi for larger 171-micron material. Weimer and Quarderer also reported that the correlation for dense-phasevoidage developedby Kmiec ( 1982) matched their results well. The Kmiec correlation for dense-phasevoidage is:

E = ~ 8Re p +

Eq. (6)

2.7 Rel.687

d

j .209

p GaO.209

.. E u i"u o "G; > .. IU "

Pressure, Atmospheres Figure

5. The effect of pressure on Umb-Umf(Varadi

& Grace).

122 Fluidization, Solids Handling, and Processing 0.9

Particle Size, microns 66

0.8

108

Dense Phase Voidage, (-)

177

0.7

Material: Activated Carbon 0.6

0.5

0.4

0.3 0

20

40

60

80

100

Pressure, Bar

Figure 6. The effect of pressure on dense phase voidage. (Weimer & Quarderer.)

Subzwari et al. (1978) also reported that bed voidage and bed expansion increased with pressure for FCC catalyst (a Group A material). However, Knowlton (1977) reported little change in either bed density or bed expansion at Ucf with increasing pressure over the range of 1 to 70 atmospheres for large, -2000/+74 micron material. The fluidized-bed voidage at minimum bubbling conditions (εmb) also has been shown to increase with increasing temperature (Fig. 7) and increasing pressure (Fig. 8).

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123

Gas: Air Material: 103-micron FCC Catalyst

0.6

emb

Foscolo and Gibilaro Model

0.5

Data

0.4

0

200

400

600

800

1000

o

Temperature, C

Figure 7. The effect of temperature on ε mb . (Rapagna et al., 1994; Yates, 1995.)

Gas: CF 4 Material: Synclyst

e mb

0.8

Data

0.7

Foscolo and Gibilaro Model

0

20

40

60

80

Pressure, bar

Figure 8. The effect of pressure on ε mb . (Crowther & Whitehead, Fluidization, Davidson and Keairns, eds., 65–70, 1978.)

24-Sep-2001 JMR


Bubbles in Fluidized Beds

Hydrodynamically, fluidized beds are considered to be stable when they are not bubbling and unstable when they are bubbling. Several researchers (Knowlton, 1977; Hoffman and Yates, 1986; Guedes de Carvalho et al., 1978) have reported that fluidized beds become “smoother” at elevated pressures (i.e., have smaller bubbles) and, therefore, are more stable at high pressures. There are generally two approaches as to what causes instability in fluidized beds. Rietema and co-workers (Rietema et al. 1993) forwarded the theory that the stability of the bed depends on the level of interparticle forces in the bed. However, Foscolo and Gibilaro (1984) have proposed that hydrodynamics determines whether a fluidized bed is stable. With the Interparticle Force Theory, interparticle forces (van der Waals, etc.) are what cause the bed to be “elastic.” Bed elasticity is characterized by an Elasticity Modulus, Mi. The criterion which determines when the fluidized bed starts to bubble is determined by the relative magnitudes of the two sides of Eq. (7). 3

Eq. (7)

4

ρp d p g 2

2

µ M mb

 150 (1 − ε )  mb  =  2 ε (2 − 2ε )  mb mb 

2

If the LHS of Eq. (7) is < the RHS: The bed is STABLE If the LHS of Eq. (7) is > the RHS: The bed will BUBBLE

In this theory, increasing pressure causes gas to be absorbed onto the surface of the particles. This results in an increase in Mi, and, by Eq. (7), an increase in the stability of the fluidized bed. The Hydrodynamic Theory of fluidized bed stability was proposed by Foscolo and Gibilaro who adapted the stability principle of Wallis. They postulated that a fluidized bed is composed of two interpenetrating fluids. One fluid is the gas phase, and the solids phase is also considered as a continuous fluid phase. In this theory, voidage disturbances in the bed propagate as dynamic and kinetic waves. The stability of the fluidized bed depends upon the relative velocities of these two waves. The velocities of the kinetic wave (uε ) and the dynamic wave (ue) are:

24-Sep-2001 JMR

Pressure and Temperature Effects Eq. (8)

uε = nut (1 − ε )ε

Eq. (9)

ue =

[3.2 gd

p

125

n−1

(1 − ε )(ρ p − ρ g )]

ρp

where n is the Richardson and Zaki exponent. If ue > uε , the bed is STABLE If ue > uε , the bed will BUBBLE Equating ue and uε and manipulating the resulting expression leads to the following stability criterion: Term 1 (T1)

Eq. (10)

(

 gd ρ − ρ g  p p 2  ut ρp 

Term 2 (T2)

)  

0.5 n −1 −0.56n(1 − ε mb ) ε mb = C 0.5

If C > 0 (if Term 1 is > Term 2), the bed is STABLE If C < 0 (If Term 1 is < Term 2), the bed BUBBLES The Hydrodynamic Theory of Foscolo and Gibilaro has been shown to predict the increase in εmb with temperature and pressure very well. This is shown in Figs. 7 and 8, respectively, for the data of Rapagna (1994) and Crowther et al., (1978). Jacob and Weimer (1987) also reported that the Foscolo and Gibilaro theory successfully predicted the increase in εmb with increasing pressure.

1.4

Bubble Size and Frequency

Temperature and pressure also interact with particle size to affect bubble size and frequency in fluidized beds. Information on the effect of temperature on bubble size in the literature is somewhat inconsistent. However, the information that does exist suggests that bubble size decreases slightly with temperature for Group A materials (Geldart and

24-Sep-2001 JMR

126 Fluidization, Solids Handling, and Processing Kapoor, 1976; Kai and Furusaki, 1985; Yoshida et al., 1976). Although less information exists for larger particle sizes, bubble size appears to not change with temperature for Group B materials (Sishtla et al., 1986; Wittman et al., 1981), and to increase with temperature for Group D materials (Sittiphong et al. 1981). Workers generally report that bubble frequency increases with temperature (Mii et al., 1973; Otake et al., 1975; Yoshida et al., 1974). There is an initial rapid increase in frequency with temperature near ambient, which then tapers off at higher temperatures. As reported above, investigators have reported that fluidization appears smoother when beds are operated at high pressure. They also report that there is a “slow motion” quality about the bed. This behavior has been attributed to a decrease in the bubble size with pressure. Indeed, there is ample experimental evidence (Barreto et al., 1984; Chan et al., 1987; Rowe et al., 1984; Weimer and Quarderer, 1983) which shows that increasing pressure causes bubble size to decrease in Group A materials. The same evidence shows that the pressure effect on bubble size decreases as the particle size increases, and that pressure does not affect bubble size significantly for large Group B, and Group D materials. The effect of pressure on bubble size for 66 micron and 171 micron material is shown in Fig. 9 from Weimer and Quarderer. As can be seen from the figure, pressure has a significant effect only for the smaller particles. Although there is some disagreement among researchers in this area, it is generally believed that bubbles divide by splitting from the roof due to Taylor instabilities (Clift and Grace, 1972), and that the smaller bubbles observed in high-pressure beds of Group A solids are a result of the increased expansion of the dense phase with pressure. This results in an effective decrease in the viscosity of the dense phase. Taylor instability (an instability in the roof of the bubble/dense-phase interface) increases with a decrease in the viscosity of the dense phase. This results in a collapse of the bubble roof more frequently at high pressures. King and Harrison (1980) used x-rays to obtain images of bubble breakup and found that bubble breakup occurred because of “fingers” of material penetrating the bubble from the roof. This is schematically depicted in Fig. 10. Experimental observations show that the dense-phase viscosity for small Group A particles decreases significantly with pressure (King and Harrison et al., 1980; May and Russell, 1953) as shown in Fig. 11. However, the dense-phase viscosity of Group B and Group D particles

24-Sep-2001 JMR


127

does not change with pressure. This supports the theory that bubbles are smaller in high-pressure beds because of Taylor instability, and fits the experimental evidence that bubbles do not change significantly with pressure in beds of Group B and Group D particles. Rowe et al. (1984) also showed that at extremely high pressures, bubbles become so small that they cannot be distinguished from the dense phase.

6

5

s (J L."' -Q) Q) s (G c Q) ::0: ..Q ~ m

4

3

2

0 0

1

3

2

u

-Umf,

4

5

6

cm/s

Figure 9. The effect ofpressure and particle size on bubble diameter. (Weimer and Quarderer.)

128 Fluidization, Solids Handling, and Processing

Particles Penetrate Into Bubble Roof Due to Taylor Instability

When Penetration Completely Pierces Bubble - Bubble Splits

Figure 10. Schematic drawing of bubbles splitting from the roof.

0.005

2

Kinematic Viscosity, m/s

0.002

475 microns

0.001 0.0005

0.0002

101 microns

0.0001 64 microns

0.00005

Material: Ballotini Gas Nitrogen Nitrogen Carbon Dioxide

0.00002 0.00001

0

5

10

15

20

25

Pressure, bar

Figure 11. Kinematic viscosity versus pressure. (King & Harrison.)

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129

Pressure also appears to cause bubble frequency to increase. This has been reported by both Rowe et al., (1984) and Chan et al., (1987). Rowe et al., (1984) also reported that bubbles were flatter at elevated pressures.

1.5

Bed-to-Surface Heat Transfer Coefficient

In general, gas-to-particle or particle-to-gas heat transfer is not limiting in fluidized beds (Botterill, 1986). Therefore, bed-to-surface heat transfer coefficients are generally limiting, and are of most interest. The overall heat transfer coefficient (h) can be viewed as the sum of the particle convective heat transfer coefficient (hpc), the gas convective heat transfer coefficient (hgc), and the radiant heat transfer coefficient (hr). Eq. (11)

h = hpc + hgc + hr

The radiant heat transfer coefficient becomes important above about 600°C, but is difficult to predict. Baskakov et al. (1973) report that depending on particle size, hr increases from approximately 8% to 12% of the overall heat transfer coefficient at 600°C, to 20 to 33% of h at 800°C. Therefore, for beds operating below 600°C Eq. (12)

h = hpc + hgc

Botterill et al. (1982) measured the overall heat transfer coefficient as a function of particle size for sand at three different conditions: 20°C and ambient pressure, 20°C and 6 atmospheres, and 600°C and ambient pressure. They found that there was a significant increase in h with pressure for Group D particles, but the pressure effect decreased as particle size decreased. At the boundary between Groups A and B, the increase of h with pressure was very small. The effect of pressure on the heat transfer coefficient is influenced primarily by hgc (Botterill and Desai, 1972; Xavier et al., 1980). This component of h transfers heat from the interstitial gas flow in the dense phase of the fluidized bed to the heat transfer surface. For Group A and small Group B particles, the interstitial gas flow in the dense phase can be assumed to be approximately equal to Umf /εd . Umf is extremely small for

24-Sep-2001 JMR

130

Fluidization,

Solids

Handling,

and Processing

these relatively small materials, and, therefore, so is hgc. For these materials, the dominant heat transfer component is hpc.The gas convective heat transfer, hgc' only becomes significant for large Group B or for Group D materials where the gas interstitial flow becomes significant. For these materials, increasing system pressure causes hgcto increase because of the increased heat transfer from the gas to a surface due to higher gas densities. The effect of pressure on the overall heat transfer coefficient for Group B particles is shown in Fig. 12.

Material: Glass Beads Gas: N2

500

.~ "' E

.

400

~ ~ 8 .r. + "0 "'

8 .c

300

200

100

0

0

0.06

0.12

0.18

0.24

0.3

0.36

Superficial Gas Velocity, Ut m/s

Figure 12. The effect of pressureon the overall heat transfer coefficient for group B powders. (Xavier & Davidson.)

Increasing system temperature causes hgc to decrease slightly because increasing temperature causes gas density to decrease. The thermal conductivity of the gas also increases with temperature. This causeshpcto increase becausethe solids are more effective in transferring heat to a surface. Because hgcdominates for large particles, the overall heat transfer coefficient decreaseswith increasing temperature. For small particles where hocdominates, h increases with increasing temperature.


131

Increasing temperature has a large effect on h for small particles near and below the Group A/B boundary. Increasing temperature causes h to increase for these particles. The effect of temperature is less pronounced for Group B particles, and h decreases with temperature for Group D materials. The primary effect of temperature on h is due to the particle convective component of the overall heat transfer coefficient. The particle convective heat transfer coefficient, hpc, depends upon heat transfer from particle “packets” to the surface. It is influenced by factors which affect the solids circulation rate. The greatest resistance to heat transfer for particle convection is when heat has to flow between particles through the gas, which has a relatively low thermal conductivity. Heat transfer is especially low for Group A and B particles, for which the flow of interstitial gas is laminar. Therefore, because increasing the temperature of the gas increases its thermal conductivity, increasing temperature increases hpc. Increasing system pressure has little effect upon hpc except for Group A particles where pressure causes an increased expansion of the dense phase, and greater heat transfer because of the increased area available to transfer heat (Botterill, 1986).

1.6

Entrainment and Transport Disengaging Height

Entrainment from fluidized beds is also affected by temperature and pressure. Increasing system pressure increases the amount of solids carried over with the exit gas because the drag force on the particles increases at higher gas densities. May and Russell (1953) and Chan and Knowlton (1984) both found that pressure increased the entrainment rate from bubbling fluidized beds significantly. The data of Chan and Knowlton are shown in Fig. 13. Increasing gas viscosity also increases the entrainment rate from fluidized beds because the drag force on the particles increases with increasing gas viscosity. Findlay and Knowlton (1985) varied gas viscosity in their experimental system (by changing system temperature) while maintaining gas density constant (by adjusting system pressure) in order to determine the effect of gas viscosity on the entrainment rate from a fluidized bed of char and limestone. They found that increasing gas viscosity significantly increased the entrainment rate from fluidized beds as shown in Fig. 14.

24-Sep-2001 JMR


4

Material: Sand Gas: Nitrogen Diameter: 30 cm Particle Size: 200 microns

3.5

Entrainment Rate, kg/m

3

3

2.5

2

Pressure, bar 4.4 11.2 21.4 31.6

1.5

1

0.5

(Knowlton and Hirsan)

0 0

0.1

0.2

0.3

0.4

0.5

Superficial Gas Velocity, m/s

Figure 13. The effect of system pressure on entrainment. (Knowlton & Hirsan.)

24-Sep-2001 JMR


133

50

Material: 65% Lignite Char/35% Limestone Particle Size: -250 microns Column Diameter: 20 cm Gas Density: 5.8 kg/m 3

Entrainment Rate, kg/h

20

10

5

2

Viscosity, kg/m-s 1

0.5 0.05

0.065 (27 C) 0.115 (427 C) 0.146 (760 C)

0.1

0.2

0.3

0.5

Superficial Gas Velocity, m/s

Figure 14. The effect of gas viscosity on entrainment. (Findlay & Knowlton.)

24-Sep-2001 JMR

134 Fluidization, Solids Handling, and Processing Chan and Knowlton (1984) also investigated the effect of system pressure on the transport disengaging height (TDH) over a pressure range of 1 to 30 bar. They found that TDH increased linearly with pressure over this range (Fig. 15). The effect of gas viscosity on TDH has not been determined experimentally by anyone thus far. 3

Material: Sand Gas: Nitrogen Column Diameter: 30 cm Particle Size: 200 microns

2.5

Gas Velocity, m/s 2

TDH, m

0.198 0.229 0.259

1.5

1

0.5 0

10

20

30

40

Pressure, bar

Figure 15. The effect of pressure on TDH. (Knowlton & Hirsan.)

1.7

Particle Attrition at Grids

Particle attrition not due to thermal or chemical reaction effects (i.e., mechanical attrition) occurs much more rapidly in the grid region of fluidized beds than in the bulk of the bed. This is due to high-velocity gas

24-Sep-2001 JMR


135

jets at the grid which accelerate the particles and cause them to collide with others at a very high velocity. Sishtla et al. (1989) reported that the particle attrition rate at the grid is proportional to the kinetic energy of the gas jets (ρgUj2 ) issuing from the grid. They varied both system pressure and temperature while maintaining gas velocity constant, and found that increasing system pressure resulted in an increase in the attrition rate (Fig. 16), while increasing system temperature decreased the rate of particle attrition (Fig. 17). They attributed this behavior to the change in gas density produced by changes in temperature and pressure. Kono (1981) also observed the same effects of temperature and pressure on mechanical attrition.

Material: 65% Coke Breeze/35% Limestone Particle Size: 1000 microns Column Diameter: 15 cm Gas: Nitrogen Fluidizing Velocity: 0.46 Grid Type: Bubble Cap

70

Fines Generated, %

60

Pressure 50

1 bar 4 bar 12 bar

40

30

20

10

0 0

2

4

6

8

10

12

14

Time, h

Figure 16. The effect of system pressure on attrition at the grid. (Sishtla, et al.)

24-Sep-2001 JMR


Particle Attrition in Cyclones

There can also be substantial particle attrition in cyclones in fluidized-bed systems because particles are accelerated at the inlet of the cyclone and impacted against the cyclone wall. Although there is little information on particle attrition in cyclones in the literature, it has been reported (Sishtla) that increasing system pressure decreases the attrition rate in cyclones operating with coal char. The mechanism by which this occurred was not determined.

100

Material: Lignite Char 90

Gas: Nitrogen Column Diameter: 15 cm Velocity: 0.46 m/s Bed Height: 30 cm Particle Diameter: 1000 microns

Fines Generatied, %

80

70

60

50 o

Temperature, C 40

21 427 871

30

20

10

0 0

2

4

6

8

10

12

14

Time, h

Figure 17. The effect of temperature on attrition at the grid. (Sishtla et al.)

24-Sep-2001 JMR

Pressure and Temperature Effects 1.9

137

Jet Penetration

Vertical jet penetration in fluidized beds increases significantly with system pressure as shown in Fig. 18. Vertical jets fluctuate in the vertical direction between a minimum and a maximum value. Most jet penetration correlations are given in terms of the maximum jet penetration length which is designated Lmax (Hirsan et al. 1980). Over a pressure range of 1 to 50 bar with jet nozzles up to 3.8 cm in diameter, Hirsan et al. (1980) developed the following empirical correlation for the maximum jet penetration.

Eq. (13)

 ρ U L max g o = 19. 3   do ρ gd p  p

   

0. 83

 U cf   U

  

0. 54

This correlation predicts that the maximum vertical jet penetration into a fluidized bed varies with gas density to the 0.67 power, and decreases with increasing fluidizing gas velocity and increasing particle diameter. Yang (1981) incorporated the high-pressure data of Hirsan et al. (1980) to produce the following correlation for Lmax

Eq. (14)

 Lmax = 7.65 1 do  Rcf 

 ρg   ρp − ρg 

2  Uo     g do   

0.472

where Rcf is defined as:

Eq. (15)

Rcf =

(U ) (U ) cf

pressure

cf

ambient

Yates et al. (1986) varied system pressure up to 20 bar with jet nozzles of 2 and 4 mm diameter, and also found that the maximum jet penetration varied significantly with pressure. They then developed the following correlation for the maximum jet penetration:

24-Sep-2001 JMR


 Lmax = 9.77 1 do  Rcf 

Eq. (16)

 ρg   ρp − ρg 

2  Uo     g do   

0.38

which is similar in form to the correlation of Yang. Yates and Cheesman also incorporated high-temperature jet penetration data (obtained at temperatures up to 800°C) into this correlation.

Particle Size: 400 microns Gas: Nitrogen Nozzle Diameter: 2.54 cm Jet Velocity: 7.6 m/s Fluidization Velocity: U cf

Lb /do

30

20

Material

10

Particle Density, g/cc

FMC Char Sand Siderite

1.2 2.6 4.0

0 0

10

20

30

40

Pressure, bar

Figure 18. The effect of pressure on jet penetration. (Knowlton and Hirsan.)

Findlay and Knowlton (1985) investigated the effect of gas viscosity on jet penetration by varying system temperature while maintaining gas density constant. This was accomplished by adjusting system pressure at

24-Sep-2001 JMR


139

each temperature studied. They concluded that viscosity had a negligible effect on jet penetration, and that temperature only affects jet penetration through its effect on gas density.

1.10 Regime Transitions The transition boundaries between various fluidization regimes (i.e., bubbling/slugging to turbulent, and turbulent to fast fluidization) are also affected by gas properties. Because the minimum fluidization velocity changes with pressure for large particles, it would be expected that the transition velocities where the bed changes from the bubbling to the turbulent regime, and where the turbulent regime changes to the fast fluidization regime, would also be affected by system pressure. Indeed this is the case. If the superficial gas velocity of a fluidized bed operating in a bubbling/slugging mode is increased, the bed is gradually transformed into a turbulent fluidized bed. The transition to turbulence occurs over a velocity range which is bounded by two unique gas velocities: Uc , the velocity at which the bed starts the transition from bubbling/slugging to turbulent, and Uk , the velocity at which the bed is fully at turbulent fluidization. Uc is often taken to be where the amplitude of the differential-pressure fluctuations in the bed are highest, and Uk is taken to be the velocity at which the amplitude of the differential-pressure fluctuations level off as indicated in Fig. 19. Recently, Marzocchella and Salatino (1996) reported that Uc and Uk can also be determined by plotting the variance of the pressure drop across the bed versus the superficial gas velocity through the bed. They conducted their tests with 175-micron glass beads fluidized by CO 2 at 40 and 60 bar pressures. They found that Uc occurred at the maximum value of the variance of the pressure drop across the bed versus superficial gas velocity curve, while Uk was determined where the variance of the pressure drop across the bed levels off (Fig. 20). This figure shows that Uc and Uk both decrease with increasing pressure. Marzocchella and Salatino also showed that power spectral densities of signals from a hot-wire anemometer inserted into the fluidized bed also can also be used to determine the packed-bed, particulate, bubbling, and turbulent regions in the fluidized bed.

24-Sep-2001 JMR


Magnitude of Pressure Drop Fluctuations

Several other researchers have investigated the effects of pressure as well as temperature on the transition from bubbling/slugging to turbulent fluidization (Cai et al., 1989; Canada and McLaughlin, 1978; Yang and Chitester, 1988). Cai et al. (1989) varied system pressure over a range of 1 to 6 atmospheres for solids with different average particle sizes. They found that Uc decreased with system pressure. Yang and Chitester (1988) varied system pressure up to 64 atmospheres and also reported that increasing pressure caused a decrease in the velocity at which the transition from bubbling to turbulent fluidization occurred.

Bubbling/ Slugging

Transition Region

Uc

Turbulent

Uk

Gas Velocity

Figure 19. Bubbling-to-turbulent transition.

24-Sep-2001 JMR


141

1.5

Variance of Pressure Drop, mbar

2

Material: Glass Beads Size: 175 Microns Gas: Carbon Dioxide Bed Diameter: 30 mm

1

Uc

0.5 40 Bar

Uk 60 Bar 0 0

10

20

30

Superficial Gas Velocity, cm/s

Figure 20. The variance of bed pressure drop versus superficial gas velocity. (Marzocchella & Salatino.)

Canada and McLaughlin (1978), working with 650-micron and 2600-micron particles over a pressure range of 1 to 10 atmospheres, found that Uk decreased significantly with pressure for both particle sizes (Fig. 21). They also reported that the ratio (Uk /Ut) remained relatively constant over this pressure range for both materials.

24-Sep-2001 JMR

142 Fluidization, Solids Handling, and Processing 6

Material: Glass Beads Column: 30-cm Square Gas: Air

Transition Velocity, U

k

, m/s

5

4

3

2600 microns 2

650 microns 1

0 0

2

4

6

8

10

12

Pressure, Atmospheres

Figure 21. Effect of system pressure on Uk . (Canada & McLaughlin.)

Cai et al. (1989) also determined Uc as a function of temperature over the range 50° to 450°C. They found that Uc increased with temperature for materials belonging to Groups A, B, and D. It appears that the decrease in gas density with increasing temperature caused Uc to increase. The boundary between the turbulent and the fast fluidization regimes has been of some dispute in the fluidization field. However, the choking velocity (Uch) appears to be a practical lower-velocity boundary for this regime (Karri and Knowlton, 1991; Takeuchi et al., 1986).

24-Sep-2001 JMR


143

Knowlton (1975), and Karri and Knowlton (1996), have also shown that the choking velocity for Group B solids decreases significantly with pressure. This is illustrated in Fig. 22 for 400-micron sand. As system pressure was increased from 1 to 31 bar, the choking velocity decreased from approximately 2.2 to 6.7 m/s for a mass flux of 210 kg/s-m2. This occurs because the carrying capacity of a gas increases as the density of the gas increases due to increased drag on the particles. This results in the pressure-drop-per-unit-length-vs.-velocity curves to be shifted to the left on the phase diagram for vertical pneumatic conveying as shown in Fig. 23.

8

Material: Sand Size: 400 microns Riser Dia.: 7.6 cm Gas: Nitrogen

1 Bar

Choking Velocity, m/s

6

4

31.6 Bar 2

0 0

50

100

150

Solids Mass Flux, kg/s-m

200

250

2

Figure 22. The variation of choking velocity with pressure. (Knowlton.)

24-Sep-2001 JMR

144

Fluidization,

Solids

Handling,

GBS Velocity, 0

5

and Processing

m/s 10

15

0.3 E ~ 0. ~

s II! 0.

...:;.

~ 2250  g 

where

Ret =

D t ρg u µg

11-Oct-2001 JMR

168 Fluidization, Solids Handling, and Processing It should be noted that in this approach, only the convective contribution to heat transfer is considered. Chen (1976) tested five different correlations, (Vreedenberg, 1958; Wender et al., 1958; Miller et al., 1951), for vertical tubes against a uniform set of experimental data obtained by Ozkaynak (1974). Figure 8 shows a comparison of the correlations and the experimental data, for a fluidized bed of 240 µm diameter glass spheres at room temperature. It is seen that there is very little agreement between the various correlations or between correlation and data. The uncertainty is of the order of 100%.

10 8 6

4

z 2'

1 1

2

4

6

8

10

Rep

Figure 8. Comparison of data with convective correlations for vertical tube in bubbling bed of 240 µm particles. (From Chen, 1974.)

11-Oct-2001 JMR

Heat Transfer in Fluidized Beds

169

Recently, Borodulya et al. (1989, 1991) described the gas film in terms of an effective conductivity divided by an effective conduction path length. Correlation of their model with experimental data gave the following equation,

Eq. (9)

ρ  Nu =0.74 Ar  s  ρg 

014 .

0.1 

0.24

 Cs    Cg 

ε 2s / 3 + 0.46RePr

ε 2s / 3 εg

where the first term represents particle convection and the second term represents gas convection. Equation (9) is claimed to be good for the following ranges of parameters: 0.1≤ d p ≤ 4.0 mm 0.1≤ p ≤ 10.00 MPa 140 ≤ Ar ≤ 1.1× 107

The correlation was tested against nine experimental sets of data with a mean-square error of 22%. Visser and Valk (1993) subsequently modified the particle convection term of the model of Borodulya et al. for low gas velocities. Their results indicated improved agreement for low velocity ranges. In 1989 Molerus and Schweinzer (1989) presented an interesting variation of this approach. These investigators reasoned that the combined convective component of heat transfer is dominated by gas convection through the matrix of particles in the vicinity of the heat transfer surface, for beds of large particles and/or at increased system pressures. Noting that this same concept would hold for packed (fixed) beds, they then developed a nondimensional correlation which is valid for fixed beds as well as for fluidized beds of Geldart’s group D particles. The convective heat transfer Nusselt number was correlated to a nondimensional group which represents the pressure drop in the bed. In fixed bed regime, the Nusselt number varies as a function of Reynolds and Euler numbers. When the gas velocity exceeds that at minimum fluidization, the bed pressure drop and the convective Nusselt number becomes constant for a given type of particles, independent of gas velocity. In the fluidized regime, the correlation of

11-Oct-2001 JMR

170 Fluidization, Solids Handling, and Processing Molerus and Schweinzer is given by the following equation for spherical (or nearly spherical) particles. Nu = 0.0247 ( Ar )0.4304 (Pr )0.33

Eq. (10)

where

Nu =

Ar =

hc d p kg

(

g d 3p ρ s − ρ g ρ gν

)

2

 Cp µ 

Pr =  k  g Figure 9 shows comparison of this model with experimental data obtained at three different pressures. The solid curve represents the relationship for fixed beds, while the dashed lines represent the behavior for fluidized beds (i.e., Eq. 10) upon exceeding minimum fluidization.

-Re

Figure 9. Comparison of data with correlation of Molerus and Schweinzer (1989).

11-Oct-2001 JMR


171

The analyses reviewed above were concerned with the contribution due to dense-particle phase conduction/convection. As noted earlier, experiential measurements do indicate that the total heat transfer coefficient increases with increasing temperature. Some significant portion of this increase may be attributed to changes in the physical properties of the gas and solid phases with temperature, resulting in altered hydrodynamics and increasing the dense phase heat transfer coefficient (h d). There are contradictory opinions regarding the significance of radiative heat transfer. Ozkaynak et al. (1983) summarized the results from eight different investigations and found significant disagreement regarding the importance of radiative heat transfer. For fluidized beds of approximately the same size particles at a common bed temperature of 1,000 oC, the radiative contribution to total heat transfer reported by three different investigations varied from 2% to 21%. Ozkaynak’s experiments showed that at bed temperatures below 500oC, the radiant contribution to total heat transfer is less than 15%. Radiant contribution increases linearly with bed temperature, becoming greater than 35% when bed temperature exceeds 800oC. The fact that this increase was not more rapid, i.e., proportional to fourth power of absolute temperature, may be attributed to the fact that convective/ conductive heat transfer also increased with temperature due to changes in physical properties of the gas. The models proposed to represent radiation transport process can be grouped into two classes. The first and simpler approach is to use some form of the Stefan-Boltzmann equation for radiant exchange between opaque gray bodies,

Eq. (11)

qr =

(

)

σ T b4 − T 4w Aw  1 − eb  + 1 +  1 − ew      Ab  eb  F  ew 

where F is the view factor from wall to bed The investigation of Ozkaynak et al. (1983), with direct experimental measurements of radiant heat flux, provides indications of the magnitude and temperature dependence of the effective bed emissivity. Taking ew and F both to be unity, these investigators calculated that the effective bed emissivity for particles like sand is in the range of 0.8 to 1.0 at bed temperatures where radiation might be significant (greater than 700oC) and is fairly insensitive to the superficial gas velocity. Thus, a very simple, approximate model for radiant heat transfer in bubbling fluidized beds would be Eq. (11), with Ab = Aw , F = 1, ew = 1, and ε w approximately equal to 0.9.

11-Oct-2001 JMR

172 Fluidization, Solids Handling, and Processing The second type of model takes a more physical representation of the radiant transport process. Generally, these models recognize that radiant photons are emitted, absorbed, and scattered by the solid particles in the fluidized bed. Bhattacharya and Harrison (1976) utilized this approach to modify the alternating layer model of Vedamurthy and Sastri (1974) by permitting radiation exchange from one layer of particles with 25 neighboring layers. The emulsion particles were treated as an absorbing and emitting medium so that radiation was attenuated exponentially with distance. A more rigorous model was presented by Chen and Chen (1981) whereby the Mickley-Fairbanks packet model was modified to include simultaneous radiative and conductive heat transfer during alternating contact of the heat transfer surface with gas bubbles and particle packets. The gas phase was taken to be transparent to thermal radiation while the particle packet was treated as a radiatively participative medium with absorption and scattering. During bubble contact, radiation was directly exchanged between parallel surfaces representing the heat transfer wall and the boundary of the bubble. During packet contact, Hamaker’s twoflux formulation of radiant transport was used to describe the absorption, scattering, and emission process within the packet dI = − (A + S)I + SJ + Aσ T 4 dy

Eq. (12)

dJ = (A + S)J − SI − Aσ T 4 dy

where

I, J A

= forward and backward radiant fluxes, respectively = volumetric absorption coefficient

S

= volumetric scattering coefficient

The one dimensional transient energy equation completed the system of equations. Eq. (13)

k pa

(

)

2 ∂T ∂ T 4 + A( I + J ) − 2 Aσ T = ρ s C ps 1 − ε s, pa 2 ∂t ∂y

These authors numerically solved the system of equations with appropriate boundary conditions to derive the time-averaged radiant and conductive heat fluxes between the fluidized bed and the heat transfer surface. Using

11-Oct-2001 JMR


173

approximate estimates for the radiant absorption and scattering coefficients, as well as for the packet residence times, this model was shown to agree with the experimental data of Baskakov et al. (1973) Brewster and Tien (1982) made a significant contribution to the solution of this problem by showing that particles act as independent point scatters as long as the interparticle clearance is greater than 1/3 wave lengths. The implication is that almost all packed or fluidized beds can be approximated as media of independent scatterers, thereby permitting direct estimation of the volumetric scattering coefficient from properties of individual particles. Cimini and Chen (1987) used a modulated flux technique to measure radiant transmission through both packed and fluidized beds of different size glass spheres. Their data showed that the independent scattering approximation gave a reasonable, though slightly low, prediction of the radiation transmission.

3.0

CIRCULATING FAST FLUIDIZATION

3.1

Hydrodynamic Characteristics

As noted earlier, increasing gas velocity for any given fluidized bed beyond the terminal velocity of bed particles leads to upward entrainment of particles out of the bed. To maintain solid concentration in the fluidized bed, an equal flux of solid particles must be injected at the bottom of the bed as makeup. Operation in this regime, with balanced injection of particles into the bed and entrainment of particles out of the bed, may be termed “fast fluidization,” FFB. Figure 10 presents an approximate map of this fast fluidization regime, in terms of a dimensionless gas velocity and dimensionless particle diameter.

 u* = u  µg  Eq. (14)

d

* p

 = d p  

(

ρ

2 g

ρs− ρ

(

g

ρ g ρs − ρ g µ

2 g

)

  g 

)g

1/ 3

   

1/ 3

= Ar

1/ 3

11-Oct-2001 JMR

174 Fluidization, Solids Handling, and Processing It is seen that for Geldart types A and B particles, fast fluidization requires superficial gas velocities approximately an order of magnitude greater than that for bubbling dense beds. In many applications of fast fluidization, the particles exiting top of the bed are captured by cyclones and recirculated for makeup injection at the bottom of the bed, hence this regime is also denoted as “circulating fluidization,” CFB.

u

II

I

I

I

I111111

Pneumatic

III

1111111

"

transport

fluidized beds

10

Ut

*

Spouted beds

~

10, Bubbling

umf

fluidized

beds

10-2

A 1 0-3

11111

r' 1

I

I

B I

I

J 1111

D 1

10

I

I

i 11111

f

102

d.* p

Figure 10. Fluidization regimes, adapted from Grace (1986) by Kunii and Levenspiel (1991); particles labeled by Geldart Classifications A, B, D.

11-Oct-2001 JMR


175

In fast fluidized beds, the particles enter the bottom of the bed with little or no vertical velocity, gaining velocity by momentum transfer from the high speed gas as they pass upward through the bed. Since particle velocity increases with elevation in the bed, mass balance requires that the average solid volumetric concentration decrease with elevation along the length of the vertical bed. Figure 11 shows representative data obtained by Herb et al. (1989). For the operating conditions represented in this figure, it is seen that the axial height required to approach fully-developed flow conditions (i.e., constant solid concentration) can be many meters, increasing with increasing particle size. The data of Hartige, Li and Verther (1986), as shown in Fig. 12 indicate that the exact rate of axial development of solid concentration varies also with solid mass flux and with bed diameter. The figures also illustrate that in the fast fluidized regime, crosssectional-averaged solid concentrations are typically less than 5%, in contrast to solid concentrations in a range of 30–40% for bubbling dense fluidization. The hydrodynamic characteristic of fast fluidized beds is complicated by the existence of significant radial (horizontal) variations in solid concentration and velocity. Radial profiles of solid volume fraction, measured by Beaude and Louge (1995) are shown in Fig. 13. It is seen that for the same Froude number, there is increasing nonuniformity of solid concentration with increasing solid mass flux. At the higher mass fluxes shown in this figure, local solid volume fractions adjacent to the bed wall approach magnitude of 30%, in contrast to solid fractions of 1–3% near the bed centerline. Herb, Tuzla and Chen (1989) and Werther (1993) noted that such radial distributions can be normalized as a general similarity profile for various operating conditions by utilizing dimensionless ratio of local solid concentration to cross-sectional-averaged solid concentration. Solid flow mass flux and velocity also vary in magnitude across the radius of fast fluidized beds. Experimental measurements obtained by Herb, Dou, Tuzla and Chen (1992) are plotted in Fig. 14. It is seen that while local solid fluxes are positive upward in the core of the bed, they can become negative downward in the region near the bed wall. The difference between core and bed regions becomes increasingly greater as total solid mass flux increases. This downward net flow of solid in the region near the bed wall has important significance for heat transfer at the wall.

11-Oct-2001 JMR


0.100 0.600 I

~

0.080 I\

~

0.060

I

o

IiJ1:!!!l 185

.6

250

A

~

0.000 I O ~ , A 88 0 2

,-

0.040

E-~ 0'

I e .6

~ I

QQ I 8

010

%[m]

" ~

O

0,000 -, 0

I

II v 0.200. .u

\

\,."

0.020

I

~

" 0.400.

t/) U) v

I

~

I 2

~

---o-o-0-0-00=:=:=:~ I

4

I 6

-~

I 8

10

z [mJ

Figure Figure 11. II. Axial Axial variation variation of of solid solid concentration concentration for for fast fast fluidized fluidized bed bed of of sand sand 2 •s. (From particles, at atUU==55m/s m/sand andGGs s==30 30kg/m kg/m2os. (From Herb, Herb, Dou, Dou, Tuzla Tuzla and and Chen, Chen, particles, 1989.) 1989.)

Figure 12. Effect of mass flux and bed diameter on solid concentration. (From Hartge, Li and Wether, 1986.)

11-Oct-2001 JMR


177 30

-I, ~ '-'

i II

O

F~131,

§ .-.

.F~135,

~

[]

~

M=16 M=15

F~129,

.F~129,

./.

20

M=12

:

A

F~130,

M=8

~

:1

A

F~129,

M=8

~ .-, -

0 ..o

...~

M=12

-!§

O C/)

...; ..

0 "...~

0 """

.." .~

"0..

.0 "" .

.Q

o

o

O i".".

~ ".. .[]

".. 0.0

0.2

0.4

.

...Q.

2r/D

0.6

0.8

/~

.~

~

10

."

-"" A...

-

t

..

0 1.0

Figure 13. Radial profiles of solid volume fraction in fast fluidized bed. Fr = U g / g d p , M = G s / ρ g U g . (From Beaude and Louge, 1995.)

(

)

Figure 14. Radial profiles of solid mass flux in fast fluidized bed. (From Herb, Dou, Tuzla and Chen, 1992.)

11-Oct-2001 JMR

178 Fluidization, Solids Handling, and Processing Another hydrodynamic complication found in fast fluidized beds is the tendency for particles to flow as aggregate clusters rather than as individual particles, Chen (1996) and Horio et al. (1988). The solid volume concentrations in such clusters also vary with radial position and operating conditions, increasing with increasing radial position and with increasing total solid flux—as indicated by the data of Soong, Tuzla and Chen (1993) shown in Fig. 15. In general, the solid concentration in clusters can be many times greater than the overall bed-averaged concentration, with significant implications for the heat transfer mechanisms.

ffI 'Q) 0.20 ~ -~ U ~

I

I

I

Vg=5.0 m/s dp=251 I"m sand Z=5.5 m 2 Gs (kg/m s) O 5B .45

0.15

§ o4J U ~ 0.10 ~ ~

0

o

0/

0-0./

Q) 8 ~ 0.05 -

-//

0 >

'C o rn

I

0.00 0.0

0.2

0.4

0.6

0.8

1.0

r/R Figure 15. Concentration of solid in clusters in fast fluidized bed. (From Soong, Tuzla and Chen, 1993.)

3.2

Heat Transfer

General Characteristics. Energy addition or extraction from fast fluidized beds are commonly accomplished through vertical heat transfer surfaces in the form of membrane walls or submerged vertical tubes. Horizontal tubes or tube bundles are almost never used due to concern with

11-Oct-2001 JMR


179

erosion in the fast-flowing two-phase flow. Thus the designer is concerned with effective heat transfer coefficients on vertical surfaces which are exposed to the fast-fluidized media. Operational parameters that are known to affect the magnitude of the heat transfer coefficient include: gas velocity, solids flow rate (mass flux), mean particle diameter, axial location of the heat transfer surface along height of the fluidized bed, radial location of the heat transfer surface across the bed, and physical dimension of the heat transfer surface. Some of the generic characteristics were indicated by the early data of Kiang et al. (1976), as shown in Fig. 16. General characteristics may be noted as:

• Magnitude of the heat transfer coefficient generally decreases with elevation along the bed. • At low elevations the heat transfer coefficient generally decreases with increasing gas velocity. • At higher elevations the heat transfer coefficient increases with increasing gas velocity. • Magnitude of the heat transfer coefficient is higher than that for equivalent air convection, but lower than that found in dense bubbling beds. These characteristics point to a correlation of the heat transfer coefficient with local solid concentration. As indicated in Fig. 11, solid concentrations tend to be higher in the bottom region of fast fluidized beds; hence the higher heat transfer coefficients in that region. Increasing gas velocity increases the upward acceleration of solid particles, causing a decrease of solid concentration at the bottom and increasing concentrations of solid at the top regions—with corresponding changes in the local heat transfer coefficients. Grace (1985) illustrates this correlation characteristic by plotting heat transfer coefficient versus the suspension density of the solid-gas mixture, using data of Fraley et al. (1983). As shown in Fig. 17, such a plot graphically illustrates the increasing heat transfer coefficient with increasing solid concentration. The various lines in Fig. 17 are correlations that attempt to capture this dependence, generally of the form,

( )

b

Eq. (15)

h= a • ρ susp

where

a = empirical function of physical properties b = empirical exponent

11-Oct-2001 JMR


Figure 16. General heat transfer characteristics of fast fluidized bed. (From Kiang et al., 1976.)

11-Oct-2001 JMR


100

500

200 SUSPENSION

181

DENSITY.

kg/m3

Figure 17. Dependence of heat transfer coefficient on solid concentration. (Data of Fraley et al., 1983, from Grace, 1985).

Parametric Effects. As indicated above, a number of parameters including properties of the gas and particles, flow rates of each phase, geometric location within the fluidized bed, operational parameters such as pressure and temperature, and the geometry of the heat transfer surface all are known to affect the magnitude of the heat transfer coefficient in fast fluidized beds. Experimental data collected at various research institutions and universities provide information on these parametric effects. Two of the parametric effects have already been touched upon above, as illustrated in Fig. 16. Examination of that figure shows that at axial positions greater than 10 diameters above the bottom of the bed, heat transfer coefficients generally increase with increasing gas velocity. It appears that for a given axial position, the coefficient would increase to a maximum value before

11-Oct-2001 JMR

182 Fluidization, Solids Handling, and Processing reaching a plateau or decreasing slowly with further increases in gas velocity. The data of Fig. 16 also illustrates a second parametric effect — that for a given gas velocity, the heat transfer coefficient generally decreases with axial elevation, the effect being strongest at relatively low superficial gas velocities. Furchi et al. (1988) measured heat transfer coefficients at the wall of a 7.2 cm diameter fast fluidized bed, operating with glass particles of various diameters. Their results, shown in Fig. 18 indicates significant parametric effects of solid mass flux and of particle diameter. For a given size of particles, it is seen that there is a marked increase in the heat transfer coefficient as the solid mass flux (Gs /Gg) increases. This trend is most obvious for the smaller size particles (d p of 109 µm), becoming less obvious with particle diameters greater than 200 µm. For a given gas velocity and solid mass flux, the data in Fig. 18 indicate significant increases in heat transfer coefficients as diameter of the fluidized particles are reduced; at a Gs/Gg of 10, the effective heat transfer coefficient is increased by approximately 300% as the particle diameter decreased from 269 to 109 µm. The interaction of parametric effects of solid mass flux and axial location is illustrated by the data of Dou et al. (1991), shown in Fig. 19. These authors measured the heat transfer coefficient on the surface of a vertical tube suspended within the fast fluidized bed at different elevations. The data of Fig. 19 show that for a given size particle, at a given superficial gas velocity, the heat transfer coefficient consistently decreases with elevation along the bed for any given solid mass flux Gs. At a given elevation position, the heat transfer coefficient consistently increases with increasing solid mass flux; at the highest elevation of 6.5 m, where hydrodynamic conditions are most likely to be fully developed, it is seen that the heat transfer coefficient increases by approximately 50% as Gs increased from 30 to 50 kg/m2s. The experiments of Dou et al. (1991) also indicate that the heat transfer coefficient varied with radial position across the bed, even for a given crosssectional-averaged suspension density. Their data, as shown in Fig. 20, clearly indicate that the heat transfer coefficient at the bed wall is significantly higher than that for vertical surfaces at the centerline of the bed, over the entire range of suspension densities tested. Almost certainly, this parametric effect can be attributed to radial variations in local solid concentration which tends to be high at the bed wall and low at the bed centerline.

11-Oct-2001 JMR

~~'..01~

rl

&

...

& 01 0101 01 01

e

:J.,CQCCCC ~CI/CI/CI/NN ...,

.

~tI'I~'..~N ~

~~'..01~

E ~~CDCD~~ ~ OIOIOIOI ,...a.Mr-or-o,.4

. ... E

~

xx**+

-w)~,.:mN ~

E "'~~~~~ -00000 ..: r-o-r-o-

. E ~~,..~~ -~~j..:~N o

~

Oo<JC..A

O

[:1

. 0 W)

~

(Jo

I>

~ "f

x

~

[:]

~

~

x

*

*

q

--

~

"t x --

orO K)

y-

.

I W)

" 1:r 0 "' ~

~ ...

. 004

11'1 -

N

I:

0 ai

ID

too

ID

VI

.. I')

IN

I...

'0

183

11-Oct-2001 JMR

Figure 18. Parametric effects of solid flux and particle diameter on heat transfer in fast fluidized beds. (From Furchi et al., 1988).


cO

o 0 rot

tW/1\\)

Gs /Gg


,u o N s -~ ~ rz.1 O u

~ rz.1 tz. (/) ~ ~ E-o E-o < rz.1 :x:

Figure 19. Interactive effects of solid flux and axial location on heat transfer in fast fluidized bed. (From Dou, Herb, Tuzla and Chen, 1991.)

Figure 20. Variation of heat transfer coefficient with radial position across fast fluidized bed. (From Dou, Herb, Tuzla and Chen, 1991.)

11-Oct-2001 JMR


185

The data of Fig. 20 also illustrates one of the most significant parametric effects, that of the solid concentration in the particle/gas suspension. Many of the parametric effects mentioned above (e.g., due to elevation, gas velocity, solid mass flux) can be attributed to their individual impact on solid concentration in the fast fluidized bed. Thus increasing gas velocity in order to increase particle entrainment to upper regions of the bed, or increasing solid mass flux to increase solid holdup within the bed, lead to higher solid concentrations and greater suspension densities—which correlates with higher heat transfer coefficients. This is illustrated by the data of Dou et al. in Fig. 20, which shows that for a given radial position (centerline or wall) the heat transfer coefficient correlates with cross-sectional-averaged suspension density. The increase of heat transfer coefficient with increasing average density is seen to be much more pronounced at the vessel wall than at the surface of a tube submerged at the bed centerline. It should be noted that these results include data obtained over a range of axial elevations, indicating that this latter parameter is secondary, influencing the heat transfer coefficient only through the effect of axial variations in suspension density. The data of Fig. 20 also point out an interesting phenomenon—while the heat transfer coefficients at bed wall and bed centerline both correlate with suspension density, their correlations are quantitatively different. This strongly suggests that the cross-sectional solid concentration is an important, but not primary parameter. Dou et al. speculated that the difference may be attributed to variations in the local solid concentration across the diameter of the fast fluidized bed. They show that when the cross-sectional averaged density is modified by an empirical radial distribution to obtain local suspension densities, the heat transfer coefficient indeed than correlates as a single function with local suspension density. This is shown in Fig. 21 where the two sets of data for different radial positions now correlate as a single function with local mixture density. The conclusion is that the convective heat transfer coefficient for surfaces in a fast fluidized bed is determined primarily by the local two-phase mixture density (solid concentration) at the location of that surface, for any given type of particle. The early observed parametric effects of elevation, gas velocity, solid mass flux, and radial position are all secondary to this primary functional dependence. The parametric effect of system pressure on the heat transfer coefficient was studied by Wirth (1995). They obtained experimental measurements of the heat transfer Nusselt number for fast fluidized beds

11-Oct-2001 JMR

186 Fluidization, Solids Handling, and Processing of various solid particles over a pressure range of 1 to 5 bars. Some of their results are shown in Fig. 22, which plots the Nusselt number as a function of Archimedes number, with solid concentration as a parameter, where the Nusselt number is defined on the basis of particle diameter and gas thermal conductivity. Wirth concluded from these observations that the effect of pressure is evident only through its influence on the hydrodynamics that result in the solid concentration in the fluidized bed. For a given gas-solid system characterized by the Archimedes number, the heat transfer Nusselt number is a function of the cross-sectional averaged solids concentration. The parametric effect of bed temperature is expected to be reflected through higher thermal conductivity of gas and higher thermal radiation fluxes at higher temperatures. Basu and Nag (1996) show the combined effect (Fig. 23) which plots heat transfer coefficients as a function of bed temperature for data from four different sources. It is seen that for particles of approximately the same diameter, at a constant suspension density (solid concentration), the heat transfer coefficient increases by almost 300% as the bed temperatures increase from 600°C to 900°C. tJ a C\2 e ~ r:. ~ O tJ ~ ~ r:. (/) ~ ~ E-. < ~ ~

Figure 21. Significance of local suspension density for heat transfer in fast fluidized beds. (Data of Dou, Herb, Tuzla and Chen, 1991.)

11-Oct-2001 JMR

O ..-

n

; 0 -

6\ o \

nN

01 o o

N

b/ -

e .. "-0 ..., NO ea.

0 O -

o

Go vlN.1-"'v. vI c:

~'-:JO ~mOa. ~mV)a.

I

l:J

e.. e ~o~o .0.#.0 "'-0-0 -0--~ -J)'

~.. e

ro .I

"'0 -0~

~o "'~

~o "' ...

l:J

Q/ c: Q/ '-

>-

e

"' l:J

.8

l:J

~.. ~ ~

e..

" 0 .#.0 ~"'

e

I])

~"'

e "0

~.. ~.8

l:J --~

00

e~ ,,-0 N.O .-~

:l:J

~V1 :;;'g -0.::r-;.8,.;S?, ~.. ;,.8

0--0","0"'2.e..

~ ~

"'

~ e

~"' ~

~

e.. "-0 ..,.0 .-~

..

e.8 ~o ~

.. ...

"' . 0 -

-0

-

C«

~C> .--

N

~o .-

"'

~ O ..-

~o ..-

.0


0'

..J ~

tJ ..!.

'i'"

~

0.

.!?;:

~

\0-:-1

\

O

0

,

,...

e

~.B

e ~

~!-~:; "' -0

Q ~,-.~~-l:J c'D-v::::::~ ~ 0-

\

O-~~

b-~

" V

~

~~ ~~ "'

E

00

~.B -o~

0 -

..

187

11-Oct-2001 JMR

Figure 22. Heat transfer coefficients for fast fluidized beds at various pressures. (Data of Wirth, 1995.)

188 Fluidization, Solids Handling, and Processing 210 SUSPENSION ~

DENSI1Y

20 kg/mJ

190 ~ BASU

( 1990), 2roulT\

0.1 m

x

NE ~ ~ z w ~ u. u. w O u a= w u. U) z ~ .-

C KO8RO

170

O '.AoNelll.

(1986), (1987),

X A NOERSON 0.05m

150

2~1T\

2SO-3OOun\

1.53m x

et al. ( 1987) , 240u"\ X

0

x

X

130

XX X ;o

C

XA

0 110

90 o

~ J:

x ><x

0.1m

jl A

70

O

~i/

50

f~

A

~~

lJ;

..I. 300

400

.. 500

600

700

TEMPERA TURE.

800

900

1000

°C

Figure 23. Effect of bed temperature on heat transfer coefficient. (From Basu and Nag, 1996.)

Another parametric effect is the apparent dependence of the heat transfer coefficient on the physical size of the heat transfer surface. Figure 24, from Burki et al. (1993), graphically illustrates this parametric effect by showing that the effective heat transfer coefficient can vary by several hundred percent with different vertical lengths of the heat transfer surface, for circulating fluidized beds of approximately the same particle diameter and suspension density. This size effect significantly contributed to confusion in the technical community since experimental measurements by inves-

11-Oct-2001 JMR


189

tigators with different size heat transfer surfaces gave apparently contradictory results. Burki et al showed in Fig. 24 that data obtained with small heat transfer probes (of centimeter dimensions) can be fourfold greater than the coefficients measured with larger heat transfer surfaces (of order of 1–2 m). This phenomenon is likely associated with a vertical thermal development length so that small probes reflect higher heat transfer coefficients representative of not-fully-developed conditions. 250

200~Basu ~ N -E

and Nag [1987]

IOou [1990] I

150

~ -;~

~ :r:

-~ ~

100 IBurki [1992]

50 I °

I 0.2

0.4

0.6

0.8

1

IWu et ai [1989]

1.2

,1.4

. 1.6

HEAT EXCHANGE SURFACE LENGTH [m]

Figure 24. Dependence of heat transfer coefficient on size of heat transfer surfaces. D p ϑ 245 µm; Suspension Density ϑ 50 kg/m3 . (From Burki, Hirschberg, Tuzla and Chen, 1993.)

Models and Correlations. A multitude of different models and correlations have been proposed for prediction of the heat transfer coefficient at vertical surfaces in fast fluidized beds. To organize the various models in some context, it is helpful to consider the total heat transfer coefficient as comprised of convective contributions from the lean-gas

11-Oct-2001 JMR

190 Fluidization, Solids Handling, and Processing phase and the dense-particle phase plus thermal radiation, as expressed by Eq. (3). Many of the models combine individual phase convection contributions into a single convective coefficient which we can define as, Eq. (16)

h c = f l hl + (1 - f l ) hd

All correlations based on ambient temperature data where thermal radiation is negligible should be considered to represent only the convective heat transfer coefficient h c. The simplest correlations are of the form shown by Eq. (15), in attempts to recognize the strong influence of solid concentration (i.e., suspension density) on the convective heat transfer coefficient. Some examples of this type of correlation, for heat transfer at vertical wall of fast fluidized beds are: Wen and Miller (1961):

Eq. (17)

hc d p  C ps   =  kg  C pg 

0.3

 ρ susp     ρ  s  

 u2   t   gdp   

0.21

Pr g

where u t = terminal velocity of particle. Kiang et al. (1976):

hc d p Eq. (18)

kg

=

ε sa

 C ps     C pg   

b

 ρ susp     ρs   

c

where a, b, c, d, e are empirical constants. Fraley et al. (1983): Eq. (19)

(

)0.90 , SI Units

(

)0.55

h c = 2.87 ρ susp

Divilio and Boyd (1993): Eq. (20)

11-Oct-2001 JMR

h c = 23.2 ρ susp

, SI Units

 Db ρ g u g   µg 

d

  Pr e g  


191

Werdmann and Werther (1993):

Eq. (21)

 Db ρ g u g  hc d p  = 7.46 × 10-4   kg  ρg 

0.757

(ρ )

0.562

susp

It should be noted that the suspension density (ρ susp) used in above correlations represent the cross-sectional-averaged value. Since Dou et al. (1991) have shown that heat transfer coefficient is primarily dependent on local solid concentration in the vicinity of the surface (rather than crosssectional-averaged concentration), the above correlations for convective coefficient at the wall of FFBs can be justified only if there is a generic similarity profile in solid concentration relating ε s at the wall to crosssectional-averaged ε s. Further study of Eqs. (17–18) indicate that there are significant uncertainties with regard to the fundamental nature of appropriate heat transfer correlations. For example, such parameters as particle size (d p), bed diameter (Db), particle terminal velocity (u t), heat capacity of solid (Cps), and gas thermal properties (ρ g , k g , µg , Cpg) are incorporated in some correlations and not in others. Obviously, each of these empirical correlations is useful only for the range of parameters covered by the correlation’s experimental data base. In contrast to the purely empirical correlations, various authors have attempted to develop mechanistically based models for each of the contributions from particle convection, gas convection and thermal radiation reflected in Eq. (3). Of these, it is generally acknowledged that the dense phase convection (represented by h d) is the dominant contributor. The lean phase gas convection (represented by h l) is often thought to be small or negligible. The radiative contribution (represented by h r) is acknowledged to be significant for high temperature fluidized beds, at temperatures greater than 500°C. Various models proposed for each of these contributing mechanisms are briefly summarized below. The lean/gas phase convection contribution has received the least attention in the literature. Many models in fact assume it to be negligible in comparison to dense phase convection and set h l to be zero. Compared to experimental data, such an approach appears to be approximately valid for fast fluidized beds where average solid concentration is above 8% by volume. Measurements obtained by Ebert, Glicksman and Lints (1993) indicate that the lean phase convection can contribute up to 20% of total

11-Oct-2001 JMR

192 Fluidization, Solids Handling, and Processing convective heat transfer when the average solid concentration becomes less than 3%, as is often found in upper regions of fast fluidized beds. A better approach is to approximate the lean phase convection by some equation for single phase gas convection, e.g., the Dittus-Boelter correlation,

Eq. (22)

hl Db = 0.023  Db ρ g u g    µg kg  

0.8

(Prg ) 0.33

Some researchers have noted that this approach tends to underestimate the lean phase convection since solid particles dispersed in the up-flowing gas would cause enhancement of the lean phase convective heat transfer coefficient. Lints (1992) suggest that this enhancement can be partially taken into account by increasing the gas thermal conductivity by a factor of 1.1. It should also be noted that in accordance with Eq. (3), the lean phase heat transfer coefficient (hl) should only be applied to that fraction of the wall surface, or fraction of time at a given spot on the wall, which is not submerged in the dense/particle phase. This approach, therefore, requires an additional determination of the parameter fl , to be discussed below. The contribution of dense phase convection is acknowledged to be most important and has received the greatest attention from researchers. Models for prediction of h d have followed two different, competing concepts. The first type of models considers the heat transfer surface to be contacted periodically by clusters or strands of closely packed particles. Similar to the packet model for dense bubbling beds, this approach leads to a surface renewal process whereby the heat transfer occurs primarily by transient conduction between the heat transfer surface and the particle clusters during their time of residence at the surface. Equation (4), from Mickley and Fairbanks (1955), can be used here to estimate the effective cluster heat transfer coefficient by treating clusters in FFBs as the equivalent of particle packets in bubbling beds. Experimental evidence supporting this concept is reported by Dou, Herb, Tuzla and Chen (1992). They obtained simultaneous measurements of the dynamic variations of solid concentration and instantaneous heat transfer coefficient at the wall of a fast fluidized bed, and found significant correlation between the two—indicating a direct dependence of the dense phase heat transfer on cluster contact. Application of this cluster type model requires information on solid concentration in the cluster (ε s ,cl ), average residence time of

11-Oct-2001 JMR


193

cluster on the heat transfer surface (τcl ), and the effective thermal conductivity of the cluster (k cl ). Figure 15 shows sample measurements of solid concentration in clusters as reported by Soong, Tuzla and Chen (1993) at various locations within the fast fluidized bed. Other measurements of ε s ,cl have been reported by Wu (1989), Louge et al., (1990), and Lints (1993). Lints report that the average solids concentration in the cluster can be correlated with the cross-sectional-averaged solid concentration across the bed, Eq. (23)

ε s ,cl = 1.23 (ε s,b )

0.54

The average cluster residence time (τcl) could be correlated directly with operating parameters (e.g., u g , d p , Gs , H) if sufficient experimental data on hydrodynamics were available. Alternately one can assume that the clusters of finite length, Lcl , fall down over the heat transfer surface with a velocity Ucl . For any given position on the heat transfer surface, the cluster contact time would just be the ratio Lcl /Ucl . Experimental measurements of cluster velocity and length have been reported by Hartge et al., (1986), Horio et al., (1988), Lints (1992), and Soong et al. (1995). Figure 25 shows sample data reported by Soong et al. indicating that near FFB wall, cluster velocity are in the range of 1–5 m/s downward, and cluster length are in the range 0.5–4.5 cm. While there is general agreement about the range of values for these two parameters, there are inadequate data to develop general correlations. In the absence of such correlations, Lints and Glicksman (1993) suggest that an average downward velocity of 1 m/s be assumed for the clusters, and that the vertical length of the heat transfer surface be substituted for the cluster length. Clearly this will require additional research and analysis. Finally, to utilize this cluster concept in modeling the heat transfer, one needs to determine the fraction of time that a spot on the heat transfer surface is covered by clusters, which can be assumed to be equal to the fractional of wall coverage by clusters at any given time. Sample measurements of f l have been obtained by Wu (1989), Rhodes et al., (1991), Louge et al., (1990), Rhodes et al., (1991), and Dou et al. (1993). Lints and Glicksman present a combined plot of some of these data, reproduced here in Fig. 26. They suggest an approximate correlation for the fractional wall coverage in terms of the cross-section-average solids concentration, Eq. (24)

(1 - f l )

= 3.5 (ε s, b )

0.37

, at wall

11-Oct-2001 JMR

(ww)

Hl~N3l

(s /w)

AlI:)013A

'd31Snl~

CJ31Snl:)


11-Oct-2001 JMR

oc '-

Figure 25. Sample data on cluster velocities and length. (From Soong, Tuzla and Chen, 1995.)


195

With the above functions and empirical correlations, it becomes possible to calculate the overall convective heat transfer coefficient h c by Eqs. (16, 4, and 22–24). Figure 26 shows a plot presented by Lints and Glicksman which compares predictions by this method with experimental data from several different sources. Reasonably good agreement is obtained over a range of bed densities corresponding to approximately 0.5 to 3% volumetric solid concentration.

~

~

I

N

E ~ c Q) .(3 ~ Q) 0 (,) ... Q) ~ c e -

200

100

m Q) .c

~

0 0

10 Cross-section

20 average

~ bed

40 density

50

(kg/m3)

Figure 26. Comparison of predictions from cluster-based model to heat transfer data. (From Lints and Glicksman, 1993.)

An alternate to the concept of cluster renewal discussed above is the concept of two-phase convection. This second approach disregards the separate behavior of lean and dense phases, instead models the time average heat transfer process as if it were convective from a pseudohomogeneous particle-gas medium. Thus h g , h cl , h l , and h d are not

11-Oct-2001 JMR

196 Fluidization, Solids Handling, and Processing calculated separately, rather a pseudo-homogeneous convective coefficient h c is sought. Attention is thus focused on the “boundary layer” of the two-phase mixture at the heat transfer surface. Different approaches are used by various authors to represent the essence of the thermal transport process in this layer. The simplest form of this approach is to borrow from single phase convective correlations, and attempt to find general empirical relationships between an effective convective Nusselt number and some appropriate Reynolds number. Equations (17–21) represent examples of this approach, with their concurrent advantage of simplicity and disadvantage of limited applicability. Martin (1984) suggested a more mechanistically based approach, utilizing analogy between kinetic motion of molecules in gases and particle motion in the FFB suspension. By drawing on the kinetic theory of gases, Martin developed a model to account for thermal energy transport across the boundary layer by particle motion. The resulting Nusselt number for particle-enhanced convection is obtained as

(

hc d p = ε s (Z ) 1 - e- Νwp /CZ kg

)

Eq. (25)

Nuc ≡

where:

εs

= average solid volume fraction in bed

Z

=

Nwp

= Nusselt number for wall-particle heat transfer

wp

= Average velocity of random particle motion

ρ p C pp d p • wp 6 kg

The wall-particle Nwp represents the heat transfer process between the wall and a contacting particle. For gas-solid systems, where thermal resistance of the gas gap dominates,

Eq. (26)

11-Oct-2001 JMR

  1  - 1 N wp ~ 4 (1 + K n ) • l n 1 +  K n    

Heat Transfer in Fluidized Beds where:

Kn = Knudsen number for gas in gap

Eq. (27)

Kn =

where:

γ

4  2  k g 2 πRT/ M g  - 1 d p  γ  P 2 C pg - R/ M g

(

197

)

= accommodation constant (≈ - 0.9 for air at ambient temperature)

Using Maxwell’s distribution for particle velocity, Martin obtained the following expression for the average random particle velocity as a function of particle size and concentration,

Eq. (28)

wp

 gd p ε s ,mf − ε s =  •  5ε s , mf εs 

   

1

2

The parameter C in Eq. (25) is a dimensionless parameter inversely proportional to the average residence time of single particles on the heat transfer surface. It is suggested that this parameter be treated as an empirical constant to be determined by comparison with actual data in fast fluidized beds. The lower two dash lines in Fig. 17 represent predictions by Martin’s model, with C taken as 2.0 and 2.6. It is seen that an appropriate adjustment of this constant would achieve reasonable agreement between prediction and data. An alternate modeling approach is that suggested by Molerus (1993) and Wirth (1993, 1995). These researchers reasoned that the wall-tosuspension heat transfer in FFBs depends on the fluid mechanics immediately near the wall. They further conclude from experimental investigations that there is a thin layer near the wall of low solid concentration, providing significant thermal resistance between the wall and falling strands of relatively high solid concentrations. Consequently Wirth and Molerus sought to describe the Nusselt number for heat transfer between the wall and the two-phase suspension in terms of dimensionless numbers which characterize the gas-particle flow near the wall. They report that a dimensionless pressure-drop number which measures bed-averaged solid concentration, and the Archimedes number for the particle/gas system are sufficient to

11-Oct-2001 JMR

198 Fluidization, Solids Handling, and Processing correlate the convective heat transfer coefficient over wide ranges of fluid properties and operating conditions. At low Archimedes numbers, heat transfer occurs primarily by gas conduction. At higher Archimedes numbers, gaseous conduction and convection both contribute to heat transfer. The correlation suggested by Wirth (1995) simply assumes that conduction and convection are additive, resulting in the following correlation.

Eq. (29)

where:

hc d p kg

  ∆P/ ∆H  = 2.85  g  ( )  ρ s ρ g ε s , mf 

(

)

0.5

+ 3.28 × 10 - 3 Recl Pr g

Re cl = Reynolds number for falling cluster strands

ρg d pU cl =

Ucl

µg

= velocity of falling cluster strands

Based on some physically reasoned assumptions, Wirth (1990) obtained the following relations for estimating Reynolds number of the falling strands:

2 (1 - φ )

Re cl =

Eq. (30)

[

2

Re mf + φ Re t

2 φ + (1 - φ ) ε gmf

]

(

)

  ½ 4 2 3 2  1 − 4 (1 − φ ) Re mf + (1 − φ ) φRe t Re mf  − 189 ε s ,mf φ (1 − φ ) Ar ]  

[(

2 1−φ

11-Oct-2001 JMR

)2 Re mf

+ φRe t

]

−1



+ 1




199

where:

(1 - φ )=

2.3∆P/ ∆H

(ρ - ρ )ε s

Re mf ≡

g

s,mf

g

ρ g u g ,mf d p µg

[

= 1.14 × 10 3 + 0.04 Ar Re t ≡

]

1/2

- 33.7

ρ g ut d p µg

Figure 27 shows a comparison of Wirth’s correlation with experimental data for particles of various materials and diameter, in FFBs operating at various pressures. While the log-log plot submerges numerical differences, the agreement is generally good. Models for estimating the radiative heat transfer contribution in fast fluidized beds parallel those proposed for dense bubbling beds, as reviewed above. The more mechanistically based approach of treating the particlegas medium as a radiative participative medium has been attempted by only a few researchers for FFBs. Chen, Cimini and Dou (1988) used the twoflux radiative model of Eqs. (12–13) to analyze simultaneous convection and radiation in high temperature FFBs. Turbulent convection was combined with decreet radiative fluxes in the three transport equations, including a term for volumetric heat generation. Sample results for fluidized combustor conditions predicted a significant interdependence of radiative and convective heat transfer. This conclusion was verified by the experimental measurements of Han (1992) who showed that the contribution of radiative heat transfer was strongly effected by bed-averaged solid concentration, decreasing with increasing solid mass flux and decreasing gas velocity.

11-Oct-2001 JMR

0.-

-=: c>'

-C"

I

2

~ a ...

"0 .0 e ... ...

1

" .. c[

,..m '.;oD ~,..

~ -a 0 .. .D ~ ~ o

.. ... c o

O'.D

~0 -

-nN

..~.. "O~c: a.. ~ I .DN-

i~i

. \

~ >-

o

8 N N " ..

I . 8 ~ " .. c(

\ ~

03 4~ !~4 ~

O

9 o ... " .. c(

00

4

0 «) " .. 0(


11-Oct-2001 JMR

'0

.., "-

... I

-~ -0

..

e

-1 5

The parameters in Eq. (13) are defined in Fig. 5. no holes within one Dm

Dm

Nh = no. of holes on this manifold

D head

Nm

= no. of manifolds on the main header supplied by single gas entry point

Figure 5. Manifold sparger grid showing the definitions of various parameters

Similarly, the main header pipe should be sized based on the following equation:

Eq. (14)

2  Dhead   N D2  m m

   

2

> 5

(b) In some instances, two to three different hole sizes are used on a given manifold to get better gas distribution.

24-Sep-2001 JMR

Gas Distributor and Plenum Design

219

(c) The gas velocity in the header/manifold pipe should be 5;  N d2   h h or

24-Sep-2001 JMR

Dhead > 0.41 m

2

2   Dhead   >5  2   122 × 0.025 


229

• Summary: For an orifice diameter of 0.025 m, the downwardly directed concentric-ring sparger has 2,400 nozzles placed on four concentric rings. The pitch is 0.04 m. Sometimes the holes are staggered on the sparger pipe. And also, it is a common practice to place two nozzles at a given crosssection as shown in Fig. 7. Example 2: For the conditions of Example 1 of perforated plate design, estimate the submerged jet height and particle-attrition rate in the fluidized bed. Solution: Perforated Plate Uh = 88.6 m/s; ρg,h = 0.64 kg/m 3; ρg,b = 0.5 kg/m 3; dh = 0.025 m; N = 1,824; dp = 60 mm; ρp = 1440 kg/m3; emf = 0.42 • Attrition Rate

φ N

(

(kg/m/hole ) = K a U h ρ g , h

)

2.5

π d h2 4

• The attrition-rate constant for 0–50 micron FCC catalyst (from Table) is Ka = 2.29 × 10-5 • Attrition rate per hole φ

(

= 2.29 × 10 -5 88 .6 0.64

N

)

2.5

π (0.025) 2 4

= 4.75 × 10-4 kg/min/hole

• Steady-state attrition rate (generation of 0–50 micron fines) φ = 4.75 × 10-4 × 1,824 = 0.867 kg/min or 1.25 tons/day (0.65% of bed) • Gas jet penetration depth using Merry’s correlation (Eq. 2) for horizontal jets

Lhor dh

 ρ g,hU h2  = 5.25  ρ (1 - ε ) gd mf p  p

   

0.4

ρ  g ,b  ρ  p

   

0.2

dp    d   h

0.2

24-Sep-2001 JMR

230 Fluidization, Solids Handling, and Processing   0.64 × 88.6 2  Lhor = 5.25  1440(1 - 0.42)9.8 × 65 × 10 -6   

From Eq. (1)

0.4

 0 .5     1440 

0.2

0.2

 65 × 10- 6    × 0.025 = 0.32 m    0.025 

Lup ≈ 2Lhor ≈ 2 × 0.32 ≈ 0.64 m

Example 3: For the conditions defined in Example 1 and 2, design a shroud having an ID twice that of the grid hole, i.e., Ds = 2dh = 0.05 m. Solution: Perforated Plate The minimum length of the shroud should be: Lmin =

0.05 - 0.025 = 2 tan 5.5°

0.13 m

• The gas jet velocity emanating from the shroud is

U h , s = U h,1

d  h D  s

2

2   = 88.6  0.025  = 22.2 m/s     0.05  

• Particle attrition rate will be reduced by a factor calculated from Eq. (16) 1.6

1.6  0.05  particle attrition without shrouds  Ds   = 3.0 =  = d  particle attrition with shrouds  0.025   h

Thus, adding a shroud to the grid reduces the attrition rate to 67% of the rate without a shroud, i.e., attrition rates with and without shrouds are 0.42 and 1.25 ton/day, respectively. 8.2

Polyethylene Reactor Grid Design

Example 4: Design a flat, perforated-plate grid for the polyethylene reactor schematically shown in Fig. 11, and calculate the gas jet penetration depth. Use a triangular pitch. System parameters are:

24-Sep-2001 JMR


231

Usup = 0.5 m/s; ρg,h = 19.2 kg/m3 ; ρg,b = 17 kg/m3 ; ρp = 641 kg/m3; ρB = 272 kg/m3; ∆Pgrid = 0.4∆Pbed; dh = 0.01 m; dp = 508 mm; εmf = 0.45; t = 0.019 m TO COMPRESSOR

7m

6.1 m 11o

4.6 m 3m

12.2 m FLAT PERFORATED PLATE ETHYLENE IN

Figure 11. Schematic of polyethylene reactor

Solution: Determine ∆Pbed and ∆Pgrid , ∆Pbed = g ρBLB = 9.8 × 272 × 12.2 = 32,520 Pa ∆Pgrid = 0.4∆Pbed = 13,008 Pa Determine the gas velocity through the grid hole, (trial and error)assume Cd = 0.8

U h = Cd

2∆ Pgrid ρf

= 0.8

2 × 13008

= 29.5 m/s

19.2

24-Sep-2001 JMR

232 Fluidization, Solids Handling, and Processing • Determine the volumetric flow rate of gas Q = U sup

π D2

= 0.5

π (4.6)2

4

= 8 .3 m3 /s

4

• Determine the number of grid holes required

Q

N=

Uh

8.3 1 1 = = 3582 2 π π dh 29.5 (0.01)2 4 4

• hole density Nd =

3582 π

= 215 holes/m 2

(4.6) 2

4 ·•

Determine the hole pitch

1

Lh =

·

N d sin 60°

1

=

215 sin 60°

= 0.073 m

Check the value for Cd t dh

=

0 .019

= 1.9

0 .01

L  From Fig. 3, Cd  h  d   h

0 .1

= 0.98

0 .1

 0.01   = 0.803 ≈ 0.80 (great guess) ∴ Cd = 0.98  0.073 

24-Sep-2001 JMR


233

• Gas jet penetration depth using Merry’s correlation (Eq. 2) for horizontal jets: Lhor dh

 ρ g,hU h2 = 5.25   ρ (1 - ε ) gd mf p  p

  19.2 × 29 .5 2  Lhor = 5.25   641 (1 - 0.45) 9.8 × 508 × 10 - 6   

0 .4

   

0.4

ρ   g,b   ρ   p 

 17     641 

0 .2

0.2

 dp    d  h  

0.2

 508 × 10 - 6       0.01 

0 .2

× 0.01 = 0.55 m

From Eq. (1) Lup ≈ 2Lhor ≈ 2 × 0.55 ≈ 1.1 m Coalescence factor

λ =

Lh

=

0.073

Lup / 2

= 0.13 < 1

1.1/ 2

∴ Jets coalesce. The low value of λ indicates that the bed of solids is probably suspended above the coalesced jets. Therefore, the solids rarely come into contact with the grid. This type of design reduces the chances of grid pluggage due to “sticky” polyethylene solids. • Summary: The perforated plate has 3,582 holes, each of 0.01 m diameter arranged in a triangular pitch of 0.073 m. The hole density is 215 holes/m 2.

NOTATIONS Cd

=

discharge coefficient, see Fig. 1

dh

=

grid hole diameter, m

dp

=

sauter mean particle size, m

D

=

diameter of fluid bed, m

24-Sep-2001 JMR

234 Fluidization, Solids Handling, and Processing Dhead =

diameter of the main header pipe, m

Dm

=

diameter of the manifold pipe, m

Ds

=

shroud or nozzle diameter, m

g

=

gravitational acceleration 9.8 m/s2

Hhigh =

elevation of highest grid hole for curved grid, m

Hlow =

elevation of lowest grid hole for curved grid, m

K

=

grid pressure-drop coefficient; Eq. (6) for upward gas entry; 0.1 for lateral and downward gas entry

Ka

=

attrition-rate constant, Eq. (15)

LB

=

operating bed depth, m

Ldown =

jet penetration for downwardly directed jet, m

Lh

grid hole pitch, cm

=

Lhor =

jet penetration for horizontally directed jet, m

Lmin =

minimum shroud or nozzle length, m

Ls

=

shroud or nozzle length, m

Lup

=

jet penetration for upwardly directed jet, m

N

=

number of grid holes

Nd

=

number hole density (holes per unit area of the bed), holes/m 2

Nh

=

maximum number of holes per manifold pipe section supplied by gas entry

Nm

=

number of manifolds on the main header supplied by single as entry point

Q

=

total volumetric gas flow entering the grid, m 3/s

t

=

grid thickness, m

Uh

=

velocity of gas through the grid hole, m/s

Usup =

superficial gas velocity, m/s

ρB

=

operating bed density, kg/m 3

ρg,b

=

density of gas at bed operating conditions, kg/m 3

ρg,h

=

density of gas entering the grid hole (plenum conditions), kg/m 3

ρp

=

particle density, kg/m 3

εmf

=

voidage at minimum fluidizing conditions, (-)

θ

=

included angle of a gas jet, degrees

24-Sep-2001 JMR

Gas Distributor and Plenum Design ∆Pbed =

pressure drop across the dense bed, Pa

∆Pgrid =

pressure drop across the grid, Pa

∆Ph =

pressure drop across the grid hole, Pa

φ

rate of formation of 0-dp micron fines, kg/min

=

235

REFERENCES Blake, T. R., Webb, H., and Sunderland, P. B., Chem. Eng. Sci., 45:365 (1990) Chen, L., and Weinstein, H., AIChE J., 39(12):1901 (1993) Hiby, J. W., Chem.-Ing.-Techn., 36:228 (1964) Karri, S. B. R., PSRI Research Report No. 60 (1990) Karri, S. B. R., Grid Design Chapter, PSRI Design Manual (1991) Knowlton, T. M., and Hirsan, I., Fluidization, (J. Grace, and J. Matsen,eds.), p. 315, Plenum Press (1980) Massimilla, L., Fluidization, (Davidson, et al., eds.), p. 133, Academic Press (1985) Merry, J. M. D., Trans. Instn. Chem. Engrs., 49:189 (1971) Mori, S., and Moriyama, A., Inst. Chem. Eng., 18:245 (1978) Roach, P. T., Fluid Dyn. Res., 11:197 (1993) Shakhova, N. A., Inzh. Fiz. Zh., 14(1):61 (1968) Siegel, R., AIChE J., 22:590 (1976) Sishtla, C., Findlay, J., Chan, I., and Knowlton, T. M., Fluidization VI, (J. R. Grace, L. W. Shemilt, M. A. Bergougnou, eds.), p. 581, Engineering Foundation , p. 581 (1989) Whitehead, A. B., in: Fluidization, (J. F. Davidson, and D. Harrison, eds.), p. 781, Academic Press (1971) Yang W. C., and Keairns, D. L., Ind. Eng. Chem. Fundam., 18:317 (1979) Yates, J. G., Bejcek, V., and Cheesman, D. J., Fluidization V, (K. Ostergaard, and A. Sorensen, eds.), p. 79, Engineering Foundation (1986) Zenz, F. A., and Othmer, D. F., Fluidization and Fluid-Particle Systems, p. 171, Reinhold Pub. Co. (1960) Zenz, F. A., Inst. Chem. Eng. Symp., 30:136 (1968) Zenz, F. A., and Kelleher, E. G., J. of Powder and Bulk Solids Tech., 4:13 (1980) Zuiderweg, Proc. Int. Symp. On Fluidization, (A. A. H. Drinkenburg, ed.), p. 739, Netherlands University Press (1967)

24-Sep-2001 JMR

5 Engineering and Applications of Recirculating and Jetting Fluidized Beds Wen-Ching Yang

1.0

INTRODUCTION

In a conventional fluidized bed, fluid under pressure is passed through a bed of solids via a distributor plate. At a fluid velocity beyond the minimum fluidization or minimum bubbling velocity, visible bubbles appear. The fluid thus passes through the bed in two phases, the bubble and the emulsion phases. The bubble-induced solids mixing and circulation provide the liquid-like behavior of a bed of otherwise immobile solids. The liquid-like behavior of a fluidized bed allows continuous feeding and withdrawal of bed material. The vigorous mixing of solids in the bed gives rise to a uniform bed temperature even for a highly exothermic or endothermic reaction. This leads to easier control and operation. The advantages of a fluidized bed, compared to other modes of contacting such as a packed bed, are numerous and they are described in details in standard textbooks on fluidization. The fluidized beds are widely employed in various industries for both physical and chemical operations.

236

Recirculating and Jetting Fluidized Beds

237

The conventional fluidized beds also possess some serious deficiencies. The bubbles which are responsible for many benefits of a fluidized bed represent the fluid bypassing and reduction of fluid-solids contacting. The rapid mixing of solids in the bed leads to nonuniform solids residence time distribution in the bed. The rigorous solids mixing in the bed leads to attrition of bed material and increases the bed material loss from entrainment. Thus for many industrial applications, the conventional fluidized beds have been modified to overcome those disadvantages. Those modifications, in many ways, alter substantially the operational characteristics of the fluidized beds, and also change the design and engineering of the beds. It is the intent of this chapter to document two of the non-conventional fluidized beds in details: recirculating fluidized beds with a draft tube and the jetting fluidized beds.

2.0

RECIRCULATING FLUIDIZED BEDS WITH A DRAFT TUBE

The recirculating fluidized bed with a draft tube concept is briefly illustrated in Fig. 1. In application as a coal devolatilizer, dry coal is introduced into the devolatilizer below the bottom of the draft tube through a coal feeding tube concentric with the draft tube gas supply. The coal feed and recycled char at up to 100 times the coal feed rate are mixed inside the draft tube and carried upward pneumatically in dilute phase at velocities greater than 4.6 m/s. The solids disengage in a fluidized bed above the top of the draft tube and then descend in an annular downcomer surrounding the draft tube as a packed bed at close to minimum fluidization velocity. Gas is introduced at the base of the downcomer at a rate permitting the downward flow of the solids. The recirculating solids effectively prevent agglomeration of the caking coal as it devolatilizes and passes through the plastic stage. Many other applications have also been reported and they will be discussed in Sec. 2.4, “Industrial Applications.” This concept was first called a recirculating fluidized bed by Yang and Kearins (1974). Several other names have also been used to describe the same concept: the fluid-lift solids recirculator (Buchanan and Wilson, 1965), the spouted fluid bed with a draft tube (Yang and Keairns, 1983; Hadzismajlovic et al., 1992), the internally circulating fluidized bed (Milne et al., 1992; Lee and Kim, 1992); or simply a circulating fluidized

11-Oct-2001 JMR


Gas Out , I I

Allernalive Solids Feed

i Gas and

Solids

-Solid Feed

--.Gas

Flow Flow

Figure 1. Recirculating fluidized bed concept—draft tube operated as a dilute phase pneumatic transport tube.

bed (LaNauze, 1976). To avoid further confusion, the recirculating fluidized bed with a draft tube will be employed here. The addition of a tubular insert, a draft tube, in a spouted fluid bed changes operational and design characteristics of an ordinary spouted bed. Notably, there is no limitation on the so-called maximum spoutable bed height. Theoretically, a recirculating fluidized bed with a draft tube can have any bed height desirable. The so-called minimum spouting velocity will also be less for a recirculating fluidized bed with a draft tube because the gas in the draft tube is confined and does not leak out along the spout height as in an ordinary spouted bed. There is considerably more operational and design flexibility for a recircu-

11-Oct-2001 JMR


239

lating fluidized bed with a draft tube. The downcomer region can be separately aerated. The gas distribution between the draft tube and the downcomer can be adjusted by changing the design parameters at the draft tube inlet. Because the draft tube velocity and the downcomer aeration can be individually adjusted, the solid circulation rate in the bed can be easily controlled. The fact that the solid circulation rate depends primarily on the entrainment rate at the draft tube inlet rather than along the surface of the entire spout, allows easier manipulation through adjustment of design and operational parameters to control the residence time and cycle time distribution. Stable operation over a wide range of operating conditions, a solids circulation rate up to 100 metric tons per hour, and a solids loading of 50 (weight of solids/weight of air) in the draft tube have recently been reported by Hadzismajlovic et al. (1992) in a 95.3 cm diameter bed with a 25 cm diameter draft tube using 3.6 mm polyethylene particles. Operating conditions for a recirculating fluidized bed can be flexible as well. The bed height can be lower than the draft tube top or just cover the draft tube top so that a spout can penetrate the bed as in a spouted bed. The bed height can also be substantially higher than the draft tube top so that a separate fluidized bed exists above the draft tube. Rather than operating the draft tube as a dilute-phase pneumatic transport tube, one can fluidized the solids inside the draft tube at lower velocities to induce the necessary recirculation of the solids. Several studies were conducted in this fashion (Ishida and Shirai, 1975; LaNauze, 1976; LaNauze and Davidson, 1976). The draft tube wall can also be solid or porous, although most of the studies in the literature employ a solid-wall draft tube. Claflin and Fane (1983) reported that a porous draft tube was suitable for applications in thermal disinfestation of wheat where control of particle movement and good gas/ solid contacting could be accomplished at a modest pressure drop. The concept can also be employed as liquid-solids and liquid-gas-solids contacting devices (Oguchi and Kubo, 1973). The important design parameters for a recirculating fluidized bed with a draft tube were identified by Yang and Keairns (1978a) as the gas bypassing characteristics of the distributor plate, the area ratio between the downcomer and the draft tube, the diameter ratio between the draft tube and the draft tube gas supply, the distance between the distributor plate and the draft tube inlet, and the area ratio of the draft tube gas supply and the concentric solids feeder. The design and operation of a recirculating fluidized bed with a draft tube are discussed below.

11-Oct-2001 JMR


Draft Tube Operated As A Fluidized Bed

When the draft tube is operated as a fluidized bed rather than a dilute phase pneumatic transport tube, it can be represented by a mathematical model developed by LaNauze (1976). The driving force for solids circulation in this case was found to be the density difference between draft tube and downcomer. The energy was dissipated by particle shear at the walls. The solids circulation rate was also found to be affected only by the distance between the distributor and the draft tube and not by the draft tube length or height of bed above it. The schematic for this system is shown in Fig. 2. Because of the lower velocity in the draft tube, the draft tube diameter tends to be larger compared to draft tube operated in a dilute phase pneumatic transport mode. One disadvantage of operating the draft tube as a fluidized bed is that if the draft tube diameter is too small or the draft tube is too high, the draft tube tends to operate in a slugging bed mode. In fact, the mathematical model developed by LaNauze (1976) described below assumes that the draft tube is a slugging fluidized bed.

Hmf

Figure 2. Recirculating fluidized bed concept—draft tube operated as a fluidized bed.

11-Oct-2001 JMR


241

The pressure balance for the dense phase in the downcomer in the circulating fluidized system shown in Fig. 2 can be expressed as: Eq. (1)

∆P1− 4 = ρ b gH mf (1 − ε bd ) −

τ d Sd Ad

A similar expression can be written for the pressure balance in the draft tube as: Eq. (2)

− ∆P2−3 = ρ b gH mf (1 − ε br ) +

τ r Sr Ar

Combining Eqs. (1) and (2), we have Eq. (3)

ρ b gH mf (ε br − ε bd ) =

τ d Sd

+

τ r Sr

Ad

Ar

It was experimentally confirmed, using capacitance probes, that the solids flow down the downcomer at close to minimum fluidization, thus ε bd = 0. The bubble voidage in the draft tube, ε br , was calculated on the basis of the velocity of a rising gas slug relative to its surrounding solids. The total gas superficial velocity in the draft tube, Ufr , can be derived to be Eq. (4)

U fr = U slugε br + U mf +

Vsrε mf 1 − ε mf

The slug velocity, Uslug , is defined as the rising velocity of the slug relative to the particle velocity at its nose and can be expressed as Eq. (5)

U slug = v p + 0.35 gD , and

Eq. (6)

v p = (U

fr

− U mf ) + V sr

Substituting Eq. (5) into (4), we have

Eq. (7)

ε br =

(U fr − U mf ) − V srε mf /(1 − ε mf ) (U fr − U mf ) + Vsr + 0.35 gD

The flow rate of particles in the downcomer and the draft tube are related by a mass balance as follows:

11-Oct-2001 JMR

242 Fluidization, Solids Handling, and Processing Eq. (8)

V sr ρ s Ar = V sd ρ s Ad = W sd Ad = W sr Ar

By solving Eqs. (4) and (7) simultaneously, the mass flux can be calculated provided the wall shear stress is known as a function of particle superficial volume flow rate. Botterill and Bessant (1973) have proposed several relationships for shear stress, however, these are not general. LaNauze (1976) also proposed a method to measure this shear stress experimentally. A similar application of the concept as a slugging lifter of solids was studied by Singh (1978) based on the two-phase theory of fluidization and the properties of slugs.

2.2

Draft Tube Operated As A Pneumatic Transport Tube

Most of the applications for the recirculating fluidized bed with a draft tube operate the draft tube as a dilute phase pneumatic transport tube. Hence we will discussed this system in more details. Downcomer and Draft Tube Pressure Drop. Typical experimental pressure drops across the downcomer, ∆P1-4, and the draft tube, ∆P2-3 , show that they are essentially similar. Successful design of a recirculating fluidized bed with a draft tube requires development of mathematical models for both downcomer and draft tube. Downcomer Pressure Drop. When the downcomer is less than minimally fluidized, the pressure drop can be estimated with a modified Ergun equation substituting gas-solid slip velocities for gas velocities (Yoon and Kunii, 1970), as shown in Eq. (9). Eq. (9) ∆P1− 4 =

µ (U gd + U pd )(1 − ε d ) 2 ρ f (U gd + U pd ) 2 (1 − ε d )  L  150  + 1 . 75  g c  d p 2φs 2ε d 2 d pφsε d 

When the downcomer is fluidized, the downcomer pressure drop can be calculated as in an ordinary fluidized bed as: Eq. (10)

11-Oct-2001 JMR

∆P1− 4 = L (1 − ε d ) ρ s


243

The voidage in the downcomer, ε d , can be assumed to be the same as the voidage at minimum fluidization, ε mf . The voidage at minimum fluidization can be determined in a separate fluidized bed. The agreement between the calculated and the experimental values is usually better than ±10% (Yang and Keairns, 1978a). When the downcomer is not minimally fluidized, the bed voidage depends on the amount of aeration and solid velocity. Use of the voidage at minimum fluidization is only a first approximation. Draft Tube Pressure Drop. The pressure drop across the draft tube, ∆P2-3, is usually similar to that across the downcomer, ∆P1-4, in magnitude. Thus, for a practical design basis, the total pressure drop across the draft tube and across the downcomer can be assumed to be equal. In most operating conditions, the pressure drop at the bottom section of the draft tube has a steep pressure gradient due primarily to acceleration of the solid particles from essentially zero vertical velocity. The acceleration term is especially significant when the solid circulation rate is high or when the draft tube is short. The pressure drop inside the draft tube is more complicated because it involves acceleration of solid particles from essentially zero vertical velocity. However, the model for calculating the pressure drop in vertical pneumatic conveying lines suggested by Yang (1977) can be applied. The acceleration length can be calculated from numerical integration of the following equation.

Eq. (11)

∆L =

∫

Upr 2 Upr1

U pr dU pr 3 4

C DSε r

− 4. 7

ρ f (U gr − U pr ) 2 ( ρ s − ρ f )d p

− ( gc +

f p U pr 2

)

2D

The solid friction factor, f p , can be evaluated with the equation proposed by Yang (1978).

Eq. (12)

(1 − ε r )  (Re )t   (1 − ε r ) f p = 0.0126 ε r 3  (Re ) p 

−0 .979

The lower limit of integration, Upr1 , is derived from Eq. (13)

W sr = U pr ρ s (1 − ε r )

with ε r = 0.5, and the upper limit, Upr2 , by the following equation:

11-Oct-2001 JMR


Eq. (14)

U pr = U gr − U t

2   1 + f pU pr  × ε 4. 7 r  2 g c D  

The total pressure drop in the acceleration region can be expressed as Eq. (15)

∆P2− 3 =

∫

L

ρ s (1 − ε r )dL + 0

+

∫

L 0

∫

L

2 f g ρ f U gr 2

0

f p ρs (1 − ε r )U pr 2 2gcD

dL

gcD  ρ (1 − ε r )U pr 2  dL +  s   gc  at ⋅L

If the draft tube height is less than the acceleration length, the integration of Eq. (15) is carried out through the whole length of the draft tube. If the draft tube height is larger than the acceleration length, the integration of Eq. (15) is carried out for the total acceleration length, and the extra pressure drop for the rest of the draft tube can then be included to give the total pressure drop in the draft tube. The suggested equations have been applied to actual experimental data satisfactorily (Yang and Keairns, 1976a). Gas Bypassing Phenomenon. Because of different design and operating parameters, the distribution of the total flow between the draft tube side and downcomer side can be very different. A summary of the design parameters that will affect the gas bypassing is presented graphically in Fig. 3 for the flat distributor plate design. Similar parameters apply for the conical distributor plate. The important design parameters that will affect gas bypassing are the area ratio between the downcomer and the draft tube, [(Dc2 - D 2)/D2], the diameter ratio between the draft tube and the draft tube gas supply or the diameter of solid feeding tube, (D/d D and D/d s), the distance between the distributor plate and the draft tube inlet, L, the area ratio of the draft tube gas supply and the concentric solids feeder, [(d D2 d s2)/d s2], and the design of the downcomer gas supply nozzle. In addition to the design parameter, the operating parameters will also affect gas bypassing. The relative strength of the concentric jets of the draft tube gas supply and the solids feeder determines the half angle of the combined jet, and the jet velocity determines the jet penetration. The jet velocity of the downcomer gas supply nozzles is also important if the jets are horizontal and directed toward the draft tube.

11-Oct-2001 JMR


245

Draft Tube

Downcomer

"-

Flat

Distributor

Plate

Concentric Solids Feeder

Figure 3. Important design parameters that affect gas bypassing

The gas bypassing phenomenon was studied by Stocker et al. (1989) by measuring the differential pressure drops between the draft tube and downcomer. A more rigorous investigation was conducted by Yang and Keairns (1978a) by injecting gas tracer, carbon dioxide or helium , continuously at different locations and taking gas samples from both the draft tube side and the downcomer side. A continuous gas sampling train fed gas into a Gow-Mac thermal conductivity gas analyzer and the tracer concentration was continuously recorded in a chart recorder. The actual amounts of gas passing through the draft tube and the downcomer were then obtained by solving mass conservation equations of tracer gas. The gas bypassing results obtained from tracer gas injection studies for a flat and a conical distributor plate are shown in Fig. 4. The flow ratio, FR, is defined as the total gas flow supplied through the draft tube gas supply and the concentric solids feeder divided by the total gas flow supplied through the downcomer gas supply. The X and Y are the actual amounts of gas passing up the draft tube and the downcomer, respectively, determined from the tracer gas injection studies. If FR equals X/Y, there is no gas bypassing. If FR is less than X/Y, some of the flow supplied through the downcomer gas supply passes into the draft tube. If FR is larger than X/Y, the reverse is true.

11-Oct-2001 JMR

246 Fluidization, Solids Handling, and Processing Effect of Distributor Plate Design. Except for the conical plate, there is no simple gas bypassing relationship as shown in Fig. 4. The scatter of the data reflects the unstable operating characteristics of the recirculating fluidized bed using a flat distributor plate and sand as the bed material. The scatter may also be due to the sampling technique. Sampling bombs of 75 cm3 capacity were used for the test series with the flat distributor plate. Continuous gas sampling trains were implemented in the test series with the conical plate. The major reason for data scattering, however, is the change in operating conditions at each data point to be discussed in more detail later.

Figure 4. Summary of gas bypassing data for a conical and a flat distributor plate.

11-Oct-2001 JMR


247

Effect of Draft Tube and Downcomer Area Ratio. When a draft tube of 9.55 cm I.D. (downcomer/draft tube area ratio = 7.8) was changed to a draft tube of 5 cm I.D. (downcomer/draft tube area ratio = 30) with other design parameters being the same, the gas bypassing reversed direction, as shown in Fig. 4. With the smaller draft tube (D/d D = 1), the gas bypasses from the draft tube side into the downcomer side for most experimental conditions, except for jet velocities in excess of 76 m/s at the concentric solids feeder; with the larger draft tube (D/d D = 1.9), the gas bypasses from the downcomer side into the draft tube side in most experiments. Effect of Distance between the Distributor Plate and the Draft Tube Inlet. Figure 4 clearly indicates that the gas bypassing phenomenon depends not only on the design parameters but also on the operating conditions. For the conical plate at a distance from the draft tube inlet of L = 21.7 cm, gas bypasses from the draft tube side to the downcomer side at a high flow ratio and reverses the direction at a low flow ratio. When the conical plate was moved closer to the draft tube inlet at L = 14.1 cm, the gas bypassing direction was exclusively from the downcomer side to the draft tube side. Gas Bypassing with a Conical Distributor Plate. Further gas bypassing results (Yang and Keairns, 1983) obtained for a conical distributor plate are shown in Figs. 5–7. The data are expressed with respect to the flow ratio (FR) as well. The flow ratio is similarly defined as the total gas flow supplied through the draft tube gas supply (No. 8 flow) and the concentric gas nozzle (No. 7 flow) divided by the total gas flow supplied through the downcomer gas supply (No. 3 flow). When the downcomer gas supply (No. 3 flow) is absent, the flow ratio FR is evaluated from the ratio of No. 7 flow and No. 8 flow. Then the total flow passing through the downcomer is calculated as Y/(X + Y). The X and Y are the actual amount of gas passing up the draft tube and the downcomer, respectively, determined from the tracer gas injection studies. When only the No. 3 and No. 7 flows are present, gas bypasses from the draft tube side to the downcomer side at a high flow ratio and reverses the direction at a low flow ratio at a distance of L1 = L2 = 21.7 cm between the conical plate and the draft tube inlet (see Fig. 5). The angle of the conical plate (a = 45° and 60°) does not seem to affect the gas bypassing characteristics greatly. When the conical plate was moved closer to the draft tube inlet at L1 = L2 = 14.1 cm, the gas bypassing direction was exclusively from the downcomer side to the draft tube side.

11-Oct-2001 JMR


Run ~

Distributor PosItion, l1

Jet t«IZlle PosItion, LZ

.-J)-I

ZL7cm

ZL7 cm

.-~.I

ZL7cm

2l7cm

141cm

14.1cm

.--51.i

14.1 cm

14.1 cm

o ---8l.1

29.3cm

ZL7cm

A ---0-1

29.3cm

2l7cm

~

.35

\,

5-1

\, .J)

249

BId Mlterial -Polyethylene Beads No. 3 Flow=0

A

FR=t«I. 1 Flow/t«l. 8Flow a:ti1'

25

'>= ~

3 Flow

.20

"' 5!:: ~

IS

~

'0 '\ I ;

.10

.os

\ \ \ 0

1

2

3

4

5

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7

8

FAconical distributor plates of different Figure 7. Gas bypassing characteristics of design configurations (No. 7 and No. 8 flows).

11-Oct-2001 JMR

250 Fluidization, Solids Handling, and Processing The gas bypassing characteristics for the cases with all three flows (No. 3, No. 7, and No. 8) are presented in Fig. 6. The gas bypassing is exclusively from the downcomer side to the draft tube side when L1 = L2. A more favorable gas bypassing characteristics can be created even for the cases where L1 = L2 = 14.7 cm by holding the No. 7 flow constant and increasing the No. 8 flow as shown for Run No. 58-i and 61-i in Fig. 6. The gas bypassing characteristics in those runs are similar to that obtained when L1 = L2 = 21.7 cm. With L1 = 29.3 cm and L2 = 21.7 cm, the gas bypassing is exclusively from the draft tube side to the downcomer side, the most favorable gas bypassing characteristics. This accounts for the high solid circulation rates observed during the experiments to be discussed later in “Solids Circulation Rate.” There are also interesting gas bypassing phenomena observed when only No. 7 and No. 8 flows are present and the flow specifically supplied to the downcomer (No. 3 flow) is absent. When L1 = L2, the fraction of the total flow passing through the downcomer has a maximum as shown in Fig. 7. At very low and very high flow ratio, defined in this case as FR = No. 7 flow/No. 8 flow, downward gas flow in the downcomer was observed. In these cases, the solids circulation rate depends primarily on the entrainment capability of the jets. With L1 = 29.3 cm and L2 = 21.7 cm, no maximum in the flow split was observed. The fraction of the total flow passing through the downcomer is linearly and inversely proportional to the flow ratio FR. The total flow passing through the downcomer can be substantially higher. The No. 7 flow in Run No. 81-i is approximately twice that in Run No. 83-i. A stronger sink created by the No. 7 flow in Run No. 81-i tends to draw the gas toward the draft tube. Solids Circulation Mechanisms and Solids Circulation Rate. Both solids circulation mechanisms and solids circulation rate are important aspects in designing and operating a recirculating fluidized bed with a draft tube. For commercial applications in the area of coating and encapsulation of solid particles, such as in coating of pharmaceutical tablets and in coating seeds for delayed germination and controlling the release rate of fertilizers, the particle residence time and cycle time are important considerations. The performance based on cycle time distribution analysis for coating and granulation was studied by Mann and Crosby (1973, 1975) and Mann (1983). Further discussions on this subject can be found in Ch. 6. Solids Circulation Mechanisms. Two mechanisms for solids circulation have been observed experimentally (Yang and Keairns, 1978a).

11-Oct-2001 JMR


251

High-speed movies (1000 to 1500 frames per second) taken at the inlet and the midsection of the draft tube with a sand bed revealed that solids transport inside the draft tube was not a conventional pneumatic transport, where uniform solid suspension prevailed, but a slugging-type transport. The high-speed movies taken at the outlet of a 23 m/s air jet showed that the air jet issuing from the jet nozzle supplying air to the draft tube was comprised of bubble rather than of a steady jet. The bubble grew from the mouth of the nozzle until its roof reached the draft tube; then the sudden suction from the draft tube punctured the roof. A continuous stream of dilute solids suspension passed through the roof into the draft tube. Simultaneously, another bubble was initiated. As this bubble grew, it pushed a slug of solids into the draft tube. The high-speed movies taken at the midsection of the draft tube exhibited alternate sections of dilute solid suspension and solids slug occupying the total cross section of the draft tube. A steady jet without bubbling can be maintained in a sand bed between the jet nozzle and the draft tube inlet with high jet velocities of the order of 60 m/s and without downcomer aeration. Once the downcomer is aerated, the solids circulation rate increases dramatically and the steady jet becomes a bubbling jet. Apparently, the inward-flowing solids have enough momentum to shear the gas jet periodically into bubbles. When the polyethylene beads (density = 907 kg/m3, average size = 2800 µm) and the hollow epoxy spheres (density = 210 kg/m3, average size = 2800 µm) were used as the bed material, a steady jet between the jet nozzle and the draft tube was always observed for all experiments conducted. Solids Circulation Rate. The solids circulation rate was obtained from the particle velocity measurements at the downcomer side by following visually the tracer particles at the wall with a stop watch. The data reported here by Yang and Keairns (1983) are for polyethylene beads (907 kg/m3 in density and 2800 µm in average particle size) and hollow epoxy spheres (210 kg/cm3 in density and 2800 µm in average particle size). The experiments were carried out in a semicircular transparent Plexiglas apparatus, 28.6 cm in diameter and 610 cm in height. The effect of downcomer aeration, of distance between the distributor plate and the draft tube inlet, and of the distributor plate design configuration on solid circulation rate is discussed below. For ease of presentation for materials of different densities, the solid particle velocity in the downcomer rather than the solid circulation rate is used.

11-Oct-2001 JMR

252 Fluidization, Solids Handling, and Processing Solids circulation rate was found to be strongly affected by the design configuration at the bottom of the draft tube and the downcomer due to changes in gas bypassing characteristics. This coupling effect indicates that an understanding of the gas bypassing characteristics is essential. Except for simple cases, the dependency of the gas bypassing characteristics on design and operating parameters are still not amenable to theoretical treatment as discussed earlier. A recent study by Alappat and Rane (1995) on the effects of various design and operational parameters on solids circulation rate essentially affirms the above conclusions. Effect of Downcomer Aeration. When only the central gas flows (No. 7 and No. 8 flows) were employed without downcomer aeration, the solids circulation rate depended primarily on the entrainment rate of the jets. The linear relationship for both bed materials (hollow epoxy and polyethylene) in Fig. 8 shows that the volumetric concentration of the solids inside the draft tube after acceleration (or the gas voidage) is approximately constant, independent of particle density. This can be readily realized by expressing the volumetric solid loading in the draft tube as follows:

Eq. (16)

φ=

U pr Ar (1 − ε r ) Gr

A straight line relationship between Upr and Gr as shown in Fig. 8 implies that the volumetric solid loading φ is approximately constant because Ar is constant and ε r can be assumed to be approximately constant when the downcomer is not fluidized. More than 85% of the gas supplied through the central No. 7 and No. 8 flows in those experiments ends up in the draft tube as can be seen from the gas bypassing data presented in Fig. 7. Aeration of the downcomer can also be provided with a conical distributor plate (No. 3 flow) with greatly increased solids circulation rate as shown in Fig. 8. At lower downcomer aeration, the solids circulation rate is essentially similar to that without downcomer aeration at a distributor plate location of L = 21.7 cm. At higher downcomer aeration, however, a substantial increase in solids circulation rate is realized with the same total gas flow rate. Apparently, a minimum aeration in the downcomer is required in order to increase substantially the solids circulation rate. For polyethylene beads, this critical aeration rate is at a downcomer superficial

11-Oct-2001 JMR

254 Fluidization, Solids Handling, and Processing gas velocity of 0.42 m/s as compared to a minimum fluidization velocity of 0.76 m/s. This critical aeration velocity in the downcomer where the solid circulation rate starts to increase substantially was obtained by tracer gas injection. Data for hollow epoxy spheres did not extend to enough gas flow rate to permit extraction of the critical aeration velocity. However, it is evident from Fig. 8 that much less gas flow is required for the hollow epoxy spheres. The minimum fluidization velocity for the hollow epoxy spheres is 0.35 m/s. The same kind of phenomenon was not observed when distributor plate was located closer to the draft tube inlet at L = 14.1 cm and when only No. 7 and No. 8 or No. 7 and No. 3 flows were used. When all three flow injection locations were used, substantial improvement in solids circulation rate is possible even at L = 14.1 cm as shown in Fig. 9. The critical downcomer aeration velocities (superficial velocities based on downcomer area) for the data shown in Fig. 9 were determined through tracer gas injection experiments to be 0.29 m/s at L = 21.7 cm and 0.22 m/s at L = 14.1 cm. Effect of Distributor Plate Design. Both conical distributor plates of included angles of 60° and 90° were used. They do not seem to affect the solids circulation rate as shown in Fig. 10. Proper location of the distributor plate and the gas nozzle, however, substantially increased the solids circulation rate. When the distributor plate was located at L1 = 29.3 cm and the concentric jet was located at L2 = 21.7 cm, No. 7 and No. 8 flows alone are enough to create high solids circulation rate as though the downcomer was separately aerated by the No. 3 flow (compare Figs. 8 and 10). This design configuration changes the gas bypassing characteristics sufficiently to provide enough aeration in the downcomer. The critical aeration in the downcomer required to promote high solids circulation rate shown in Fig. 10 was determined to be 0.25 m/s (a superficial velocity based on the downcomer area) through tracer gas injection studies. This design configuration has decisive advantages in that location of the central jet at L2 = 21.7 cm from the draft tube inlet minimizes the start-up problem (discussed later in the section “Start-up and Shutdown Considerations), and location of the distributor plate at L1 = 29.3 cm eliminates the necessity of supplying separate aeration to the downcomer through an additional location such as No. 3 flow. In some industrial applications, this design feature may prove to be a critical advantage.

11-Oct-2001 JMR


T*

Gasf..

SCfM

,~

~

J.~.Ll.!.

~~.Ll -:::. -.:::. a.1 R.I Ct.1 1 Do-.:::-

lClc.

lClc.

.

a.rc.

a.rc.

.

~.Jc. ~.'c. ~.Jc. ~.'c.

a.rc. a.rc. a.7c. a.rc.

. . . .

a.rc.

a.rc.

6.

..J f8 ..:

,..,

I...

--

l.S

,1..SC,I~ ~ICI ,

.

. ~

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I LO

j

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I > ~ ~ ~

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/ / /

0.,

/

/ /

/ /0 0

o.t

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2.0 SC-

z.t

1.0

15

Figure 10. Comparison of solids circulation rate at different distributor plate design configurations.

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257

Effect of Distance Between Distributor Plate and Draft Tube Inlet. As expected, the closer the distance between the distributor plate and the draft tube inlet the lower the solids circulation rate as shown in Figs. 8 and 9. This is not only because of the physical constriction created by locating the distributor plate too close to the draft tube inlet but also because of the different gas bypassing characteristics observed at different distributor plate locations as discussed earlier. When the distance between the distributor plate and the draft tube inlet becomes large, it can create start-up problems discussed in Yang et al. (1978).

2.3

Design Example for a Recirculating Fluidized Bed with a Draft Tube

From the experimental evidence, the design of a recirculating fluidized bed with a draft tube involves the specification of a number of design parameters and an understanding of the coupling effects between the design and the operating variables. A procedure is presented here for the design of a bed to give a specified solids circulation rate. This design procedure assumes that the solids and gas characteristics, feed rates, and operating temperature and pressure are given. The design parameters to be specified include the vessel diameter, draft tube diameter, draft tube height, gas distributor, and distributor position. These parameters can be specified using the solids circulation rate model, experimental data on gas bypassing, and process requirements (e.g., selection of gas velocity in the bed above the draft tube). Determination of the Gas Bypassing Characteristics of the Distributor Plate Experience indicates that a simple theoretical model to predict gas bypassing accurately that takes into account all the design and operating variables cannot be developed. Empirical correlations, however, can be obtained by conducting experiments with tracer gas injection for a given distributor plate design at different operating conditions and at different distances from the draft tube inlet. Distributor plate designs can be studied and an optimum design selected to provide the desired solids circulation rate. Design for Desired Solids Circulation Rate It is assumed that the total gas flow into the bed is known. When the operating fluidizing velocity is selected for the fluidized bed above the draft tube, the diameter of the vessel is determined. The final design decisions include selection of the draft tube diameter, the distributor plate design, the separation between the draft

11-Oct-2001 JMR

258 Fluidization, Solids Handling, and Processing tube and the distributor plate, and the draft tube height. Selection of draft tube height may be determined by other considerations, such as solids residence time, though it also affects the solids recirculation rate. A gas distributor is selected to be compatible with the process and to maintain the gas velocity in the downcomer near Umf . The draft tube diameter is then selected by using the solids circulation rate model to obtain the desired circulation rate. The results of an example calculation for a recirculating fluidized bed coal devolatilizer of 0.51 m in diameter handling coal of average size 1200 µm at 870°C and 1550 kPa are presented in Fig. 11. The calculation is based on operating the fluidized bed above the draft tube at 4 times the minimum fluidization velocity. It is also based on the selection of a distributor plate to maintain the downcomer at the minimum fluidization condition. If the two-phase theory applies, this means that the slip velocity between the gas and the particles in the downcomer must equal to the interstitial minimum fluidizing velocity as shown below.

U fd + Upd =

Eq. (17)

Umf ε mf

where Ufd and Umf are positive in the upward direction and Upd is positive in the downward direction. The design procedures are thus: Operating Temperature Operating Pressure Vessel Diameter Particle Size Particle Density Draft Tube Height

.= e ..e' 1200 "'i6 ~ c 1(XX) .!2 5 800 ~ ~ u 'u (iXJ "C ~ 400

,, Slugging Transport in:

0

c "e :9 2800 2400 2(XX)

Choking Point 1&X1

200 0I

=8700C = 1550 kPa =0.51 m = 1200 lAm =0.959/ cm = 3. 05 m

I D~tt T,ube I: 2

: I +Dilute

1200 800

Phase Transport in,Dra~ Tu~e

oi ";Q ~ c ,9 ";Q "5 u "~ u :2 "0 VI

400

,

I

O

4 6 8 10 12 14 16 18 20 Area of Downcomerl Area of Dratt Tube

Figure 11. Projections of solid circulation rate at constant total flow and changing bed geometry—results of example calculation.

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259

1. Assume a solid circulation rate per unit draft tube area, Wsr , and calculate the particle velocity in the downcomer, Upd , from the following equation Eq. (18)

Wsr = U pd (1 − ε mf ) ρ s

Ad Ar

2. Calculate Ufd from Eq. (17) and the pressure drop across the downcomer, ∆P1-4, from Eq. (9) assuming ε d = ε mf. 3. Use trial and error between Eqs. (12)–(14) to evaluate Upr (the equilibrium particle velocity in the draft tube after acceleration), ε r (the equilibrium voidage in the draft tube after acceleration), and f p (the solid friction factor). 4. Numerically integrate Eq. (11) to obtain the particle acceleration length and Eq. (15) to obtain the pressure drop across the draft tube, ∆P2-3. 5. Compare the pressure drop across the downcomer, ∆P1-4, and that across the draft tube, ∆P2-3. If they are not equal, repeat procedures 1 through 5 until ∆P1-4 ≅ ∆P2-3. The results in Fig. 11 show that the solids circulation rate increases with a decrease in downcomer/draft tube area ratio; however, the operating condition inside the draft tube will eventually approach the choking condition where a slugging-type transport prevails. In the example, this may occur at a downcomer/draft tube area ratio of about 10. Correlations for predicting the choking point in a vertical pneumatic conveying line are available (Punwani et al., 1976; Yang, 1983). At an even lower downcomer/ draft tube area ratio, the solids circulation will approach the conditions of the low-velocity circulating fluidized bed described by LaNauze (1976). His model can be used for solid circulation calculations in the low draft tube velocity regime. Start-up and Shutdown Considerations. Both cold flow experiments and actual pilot-plant experience show that, if operating conditions and design parameter are not selected carefully, start-up (initiate solid circulation) might be a problem (Yang et al., 1978). The primary design parameters that will affect the start-up are the distance between the grid and the draft tube inlet (L) and the diameter ratio between the draft tube and the draft tube gas supply nozzle and the concentric solid feeder (D/d D and D/d s). The maximum allowable distance, L, can be determined by applying the jet penetration equation suggested by Yang and Keairns (1978b)

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260 Fluidization, Solids Handling, and Processing 1/ 2

Eq. (19)

L dD

 ρ U 2j  f = 6.5  ⋅  ρ − ρ gd  f D   s

where Uj is the gas velocity issuing from the draft tube gas supply. For high-temperature and high-pressure operations, Eq. (23) to be discussed later should be used for calculating L. Another consideration is that the jet boundary at the end of jet penetration is preferably with the physical boundary of the draft tube inlet. Merry’s expression (1975) for jet half angle can be used for this purpose

Eq. (20)

 ρf d D   cot (θ ) = 10.4   ρs d p   

0 .3

or Eq. (21)

L=

(D − d D ) 2⋅ tan (θ )

The L selected for the design should be the smaller one of that estimated from Eqs. (19) and (21). A start-up technique described by Hadzismajlovic et al. (1992) is worthy of consideration if the draft tube gas supply is retractable. The draft tube gas supply nozzle can be inserted into the draft tube during start-up and shutdown. This will reduce the difficulty described here during startup. After start-up, the supply nozzle can be lower to below draft tube inlet at a predetermined height to provide the normal operation configuration. This will prevent solids to drain into the gas supply nozzle during shutdown. Of course, if the draft tube gas supply nozzle is not movable due to hostile operating conditions, the technique can not be used. The design precautions discussed above during start-up should then be followed. Multiple Draft Tubes. Studies in the past always concentrate on beds with a single draft tube. A literature survey failed to uncover any reference on operation of multiple draft tubes. Even in the area of conventional spouted beds, the references on multiple spouted beds are rare. Foong et al. (1975) reported that the multiple spouted bed was inherently unstable due to pulsation and regression of the spouts. Similar instability was also

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261

observed by Peterson (1966) who found that vertical baffles covering at least one half of the bed height were necessary to stabilize the operation. In an industrial environment where solids are processed in large vessels, multiple draft tubes may be both necessary and beneficial. Exploratory tests in a two-dimensional bed with three draft tubes were reported by Yang and Keairns (1989). A schematic of the two-dimensional test apparatus with three draft tubes is shown in Fig. 12. The two-dimensional bed is constructed with transparent Plexiglas plates in the front and aluminum plates at the back with a cross-section of 50.8 cm by 2.54 cm and 244 cm high. The three draft tubes have a cross-sectional area of 2.54 cm by 2.54 cm each and 91 cm high. The three draft tubes divide the bed into four separate downcomers. The two downcomers next to the side walls have a cross-section of 5.9 cm by 2.54 cm while the remaining two downcomers have a cross-section exactly two times, i.e., 11.8 cm by 2.54 cm. If all three draft tubes operate similarly, the bed should have three identical cells, each with a single draft tube. The distance between the draft tube inlet and the air distributor plate was maintained at a constant spacing of 5.1 cm throughout the experiments. Polyethylene beads of relatively narrow size distribution with a harmonic mean diameter of 2800 mm and a particle density of 910 kg/m3 were used as the bed material. A static bed height of 1.4 m was employed. Two different series of experiments were carried out. In one series, the three draft tube velocities were maintained essentially constant while the aeration to downcomers was varied. One of the three draft tube velocities was purposely increased to simulate possible unbalanced operation conditions in an actual industrial plant in the second series of experiments. Each experiment was characterized by solid particle velocity in each downcomer, the pressure drop across each draft tube, and the pressure drop across each downcomer. When all three draft tubes were operated at similar velocities, the pressure drops across all draft tubes and downcomers were comparable. However, solid particle velocities in outside downcomers close to the walls were substantially less due to wall effect and redistribution of downcomer aeration flow. Smooth operations under these conditions were possible. The solid particle velocities in outside downcomers can be increased by enlarging the downcomer cross-section or by increasing downcomer aeration through separate plenums to minimize wall effects.

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EXHAUST

Figure 12. Schematic of a two-dimensional test apparatus with three draft tubes.

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263

When one of the three draft tube velocities was increased to simulate upset conditions, stable operations were still possible. These upset conditions could also be detected by pressure drop differences among various draft tubes and downcomers when differences in draft tube velocities were large. For severe upset conditions, where some of the draft tubes become downcomers, pressure drop measurement alone could not distinguish the solids flow pattern inside the draft tubes. The design methodology proposed earlier for beds with a single draft tube is still applicable here for beds with multiple draft tubes.

2.4

Industrial Applications

The application of the recirculating fluidized bed with a draft tube was probably first described by Taskaev and Kozhina (1956). They utilized the bed for low temperature carbonization of coals in a 15 cm diameter column with a 2.5 cm diameter draft tube. A “seeded coal process” was later developed by Curran et al. (1973) using the same concept to smear the “liquid” raw coal undergoing the plastic transition onto the seed char and the recirculating char during low temperature pyrolysis. Westinghouse successfully demonstrated a first stage coal devolatilizer with caking coals in a pilot scale Process Development Unit employing a similar concept where the downcomer was fluidized and the jet issuing from the draft tube was immersed in a fluidized bed above the draft tube (Westinghouse, 1977). The same concept was also proposed for extending fluidized bed combustion technology for steam and power generation (Keairns, et al., 1978). The British Gas Council has also developed the concept for oil and coal gasification (Horsler and Thompson, 1968; Horsler et al., 1969). The development eventually resulted in a large-scale recirculating fluidized bed hydrogenator gasifying heavy hydrocarbon oils (Ohoka and Conway, 1973). McMahon (1972) also described a reactor design for oil gasification using a multiplicity of draft tubes. The Dynacracking process developed by Hydrocarbon Research Inc. in the 1950’s (Rakow and Calderon, 1981) for processing heavy crude oil also utilized an internal draft tube. More recently, gasification in a recirculating fluidized bed with a draft tube was described by Judd et al. (1984) and coalwater mixture combustion, by Lee and Kim (1992). Other industrial applications of the concept include that for coating tablets in the pharmaceutical industry (Wurster et al. 1965), for drying of

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264 Fluidization, Solids Handling, and Processing dilute solutions containing solids (Hadzismajlovic, 1989), and for mixing and blending (Decamps et al. 1971/1972; Matweecha, 1973; Solt, 1972; Krambrock, 1976). Both Conair Waeschle Systems and Fuller Company supply commercial blenders based on the concept. The concept was also proposed as a controllable solids feeder to a pneumatic transport tube (Decamps et al., 1971/1972; Silva et al., 1996). Although most of the experimental data reported here were obtained with large particles, Geldart Class B and D powders, it is believed that the concept can equally be applied for any fine aeratable and free-flowing solids, Geldart's Class A powders. A similar concept has also been used for liquid-solids and liquid-gassolids contacting devices (Oguchi and Kubo, 1973; Fan et al., 1984) and bioreactors (Chisti, 1989). Bubble columns fitted with draft tubes have also been employed in the chemical process industries as airlift reactors for gas-liquid contacting operations. Examples are the low-waste conversion of ethylene and chlorine to dichloroethane, biological treatment of high strength municipal and industrial effluent and bioreactors. Critical aspects of the design and operation of bubble columns with draft tubes have recently been reviewed by Chisti and Moo-Young (1993). Freedman and Davidson (1964) also carried out a fundamental analysis for gas holdup and liquid circulation in a bubble column with a draft tube. Extensive experimentation in a bubble column with a draft tube was conducted by Miyahara et al. (1986) and in-depth analysis by Siegel et al. (1986). The effects of geometrical design on performance for concentric-tube airlift reactors were studied by Merchuk et al. (1994).

3.0

JETTING FLUIDIZED BEDS

In a gas fluidized bed, the introduction of gas is usually accomplished through distributors of various designs. Any time the gas is distributed through orifices or nozzles, a jetting region appears immediately above the grid. A large fluctuation of bed density occurs in this zone, indicating extensive mixing and contacting of solids and gas. If the chemical reactions between gas and solids are fast, much of the conversion may occur in this jetting region. Another type of fluidized bed, where the jetting phenomena is an important consideration, is the spouted fluid bed, where a large portion of

11-Oct-2001 JMR


265

gas goes through a fairly large nozzle. Because of the dominant effect of this jetting action provided by the large nozzle, this type of fluidized bed can be more appropriately called a jetting fluidized bed, especially when the jet does not penetrate through the bed like that in a spouted bed. Jetting, bubbling dynamics, and solid circulation are important hydrodynamic phenomena governing the performance and operation of large-scale jetting fluidized beds. They are the focus of our discussion in this section.

3.1

Jet Penetration and Bubble Dynamics

Gas jets in fluidized beds were reviewed by Massimilla (1985). A more recent review is by Roach (1993) who also developed models to differentiate three jet flow regimes: jetting, bubbling and the transition. However, most of the data were from jets smaller than 25 mm. The discussion here will emphasize primarily large jets, up to 0.4 m in diameter, and operation at high temperatures and high pressures. The gas jets can also carry solids and are referred to as gas-solid two-phase jets in this discussion. Momentum Dissipation of a Gas-Solid Two-Phase Jet. Gas velocity profiles in a gas-solid two-phase jet inside a fluidized bed were determined at five different horizontal planes perpendicular to jet direction using a pitot tube (Yang and Keairns, 1980). The velocity profiles were integrated graphically, and gas entrainment into a jet was found to occur primarily at the base of the jet. The measured impact pressures and static pressures were converted to gas velocities using the following equation:

Eq. (22)

U jr = C

2 g c ( Pi − Po ) ρf

The coefficient C was assumed to be 1. The typical velocity profiles are shown in Figs. 13–15 for a nominal jet velocity of 62.5 m/s and solid loadings (weight of solids/weight of gas) of 0–1.52. Note that at 1.7 cm from the jet nozzle, the velocity profile is independent of solid loadings (Fig. 13). The jet at 1.7 cm from the nozzle is still in the potential core region. At greater distances from the jet nozzle, the velocity profiles become dependent on the solid loadings (Figs. 14–15), with jets of higher solid loadings penetrating deeper into the bed, as expected.

11-Oct-2001 JMR


70

Run No.

Sy~1

Solkt R.t. (kq I hrl

GSF-I

o

0

GSF-3

.SO.

.0.40

GSF-4

.119.

.0.92

GSF-5

.197.0

Solkt ~ing

( ~)

0

1.52 62. 5 m/s

NomlNI J.t V.l«ity 65

ss .2:'50 4S

3S

ZS

Back Wall \..

20 IS 10

,..

S

~

0

0

I

2

Distance

3

4

5 from

6

7 the

8 Front

9

10

II

Plate,

~ 12

13

14

cm

Figure 13. Jet velocity profiles at 1.7 cm from the jet nozzle.

11-Oct-2001 JMR


65 ~ 55 50 "" -..45 e 40 §

Run No.

Sy~1

GSF-1

o

SolklRilt

GSF-3

.SO.

GSF~

.119..

GSF-5

.197.

Ik9/hr.1

267

SOIk1lO-'ln9(~)

0

0 8

0. 40 0.92

0

I. SZ

Nominil Jet Velocity 62. S mls A ..

i

...\ "

35

~ 30 Q.) > 25 "" fa ~

Back Wall

10r ~~ Jet Nozzle ):i.j.-~-,.,-

. o .9

\..

00 ~~:1:R..~v ; .~ .7 i t 15 I~ ~+ Distance from the Front Plate. cm

Figure 14. Jet velocity profiles at 9.3 cm from the jet nozzle.

11-Oct-2001 JMR


Figure 15. Jet velocity profiles at 23.6, 33.8, and 44.5 cm from the jet nozzle.

11-Oct-2001 JMR


269

A reasonably consistent universal velocity profile is obtained by plotting (Ujr - Ujb)/(Ujm - Ujb) vs. r/r1/2 in Figs. 16–18, comparable with the Tollmien solution for a circular homogeneous jet in an infinite medium (Abramovich, 1963; Rajaratnam, 1976).

Pltot Tube lOCItion -

.0 :::>

cm From Jet Nozzle .1.7 o 9.3 A 23.6

I e ~

I. O ~

o 33.8 c 4S. S Sol~ Lo.tln9 0.40

0 -

Tollml.n' s Solution

~

0. ~ I

~ :::>

OL

,~

0

0. S

1.0

1.5 rl

2.0

2.5

r;

Figure 16. Universal velocity profile for Run GSF-3—comparison with Tollmien’s solution.

Pltot Tube L~tion .c ~ I .6 ~ -

.

I. O

-

cm From Jet Nozzle .1.7 o 9.3 .23.6 ~ 33.8 a 45.5

~

Solil L081in9 1. 5Z

-.Q ~

0.

I '~ -

, '--s:.. --' 0

al0.5

1.0

1.5

r/

Z.O

Z.5

r;


11-Oct-2001 JMR


~ -


Jet Penetration and Jet Half Angle. Jet Penetration Jetting phenomena were studied by Yang and Keairns (1978b) in a semicircular column 30 cm in diameter using hollow epoxy spheres (ρ p = 210 kg/m3) as the bed material and air as the fluidizing medium to simulate the particle/gas density ratio in actual operating conditions at 1520 kPa and 1280°K. A two-phase Froude number, defined as (ρ f Uj2)/(ρ s - ρ f)gd o and derived from both the momentum balance and the dimensionless analysis, was found to correlate well with jet penetration data. The jet penetration depth was defined similarly to that by Merry (1975) and shown in Fig. 19. Other variations were found as described by Kececioglu et al. (1984). The correlation was extended to cover the high-pressure jet penetration data of Knowlton and Hirsan (1980) at pressures up to 5300 kPa for fluidized beds of sand (ρ s = 2629 kg/m3), FMC char (ρ s = 1158 kg/m3), and siderite (ρ s = 3988 kg/m3). A subsequent analysis by Yang (1981) indicated that the twophase Froude number originally suggested could be modified slightly to account for the pressure effect. The original two-phase Froude number took into account the effect of pressure on the dynamic pressure created by the jet through the momentum flux term ρ f Uj2 but did not include a similar correction for the changes in fluidized bed properties due to pressure. The resulting correlation is shown in Eq. (23).

11-Oct-2001 JMR


271

Figure 19. Definition of jet penetration depth and jet half angle.

Eq. (23)

Lj do

 1 ρf U 2j   = 7.65  ⋅ ⋅  Rcf ρs − ρ f gdo   

(

0 .472

)

where Eq. (24)

( ) / (U )

Rcf = U cf

p

cf atm

In the absence of (Ucf)p and (Ucf)atm, (Umf)p and (Umf)atm can be employed.

11-Oct-2001 JMR

272 Fluidization, Solids Handling, and Processing The limiting form of Eq. (24) at atmospheric pressure (101 kPa), where the correction factor Rcf = 1, approaches the correlation originally proposed for atmospheric condition, as shown in Eq. (25):

Eq. (25)

Lj do

 ρ U 2j  f  = 6.5  ⋅  ρ s − ρ f gdo   

(

)

0.5

= 6.5 ⋅ ( Fr ) j

0.5

Equation (25) now is only a special case of the general correlation expressed in Eq. (23). For concentric jets and gas-solid two-phase jets, the jet momentum flux term ρ f Uj2 can be evaluated as follows, as suggested by Yang et al. (1984b):

(

Mo Vo + M g V g + M sVs

Eq. (26)

ρ f U 2j =

Eq. (27)

Vs = Vg − U t

)

Ao

The solids particle velocity in the gas-solid two-phase jet can be calculated as shown in Eq. (27), assuming that the slip velocity between the gas and the solid particles equals the terminal velocity of a single particle. It should be noted that calculation of jet momentum flux by Eq. (26) for concentric jets and for gas-solid two-phase jets is only an approximation. It involves an implicit assumption that the momentum transfer between the concentric jets is very fast, essentially complete at the jet nozzle. This assumption seems to work out fine. No further refinement is necessary at this time. For a high velocity ratio between the concentric jets, some modification may be necessary. The good agreement obtained for all data using the modified Froude number signifies the physical significance of the parameter. In fact, the dependence of jet penetration on the two-phase Froude number can be derived theoretically from the buoyancy theory following that of Turner (1973). The momentum flux at the orifice, M, is given by Eq. (28)

11-Oct-2001 JMR

M=

π 4

d o2U 2j


273

The buoyancy flux at the orifice, B, can be expressed as

Eq. (29)

B =

ρ  d o2U j g  b − 1 ρ  4  f 

π

The characteristic length scale, Lj, taken here to be the jet penetration depth is shown by Turner (1973) for a buoyancy jet to be Lj

Eq. (30)

do

≈

M 3/ 4 B

1/ 2

do

= (Fr ) j

0 .5

Equation (30) is essentially similar to Eq. (25). The experimental data included in the correlation, however, are mostly from jets of less than 6 cm in diameter. The question often asked is whether the developed correlation is applicable for scaleup to a much larger jet nozzle. The size effect will have to be verified with operational data from a jet nozzle considerably larger. Jet penetration data were obtained for a 0.4 m diameter compound jet in a 3-m diameter cold model. The experimental jet penetration data were compared with four existing correlations (Yang et al., 1995). These correlations were selected because they are all dimensionless in form and are often quoted in the literature. They were usually derived through either dimensionless analysis or semitheoretical argument. Pure empirical but dimensional correlations are not included because the dimensional empirical correlations based on small jet nozzle data base have no value for scaleup purpose. Correlations by Merry (1975), Wen et al. (1982), and Hirsan et al. (1980) predicted substantially higher jet penetration depth than that observed, probably because they included a dimensionless factor, do /d p. Inclusion of this dimensionless group can exaggerate the jet penetration depth when the jet nozzle diameter, d o , is large. Better agreement was obtained by using the correlation proposed by Yang (1981), i.e., Eq. (23). Jet Half Angle. Determination of jet half-angle is shown also in Fig. 19. The jet half-angle can thus be calculated from the experimentally measured bubble size and jet penetration depth as follows:

Eq. (31)

D −d  o  θ = tan −1  B  2L  j  

11-Oct-2001 JMR

274 Fluidization, Solids Handling, and Processing Experimentally observed jet half-angle range from 8° to 12° for the experimental data mentioned above. These compare to 10° suggested by Anagbo (1980) for a bubbling jet in liquid. Bubble Dynamics. To adequately describe the jet, the bubble size generated by the jet needs to be studied. A substantial amount of gas leaks from the bubble, to the emulsion phase during bubble formation stage, particularly when the bed is less than minimally fluidized. A model developed on the basis of this mechanism predicted the experimental bubble diameter well when the experimental bubble frequency was used as an input. The experimentally observed bubble frequency is smaller by a factor of 3 to 5 than that calculated from the Davidson and Harrison model (1963), which assumed no net gas interchange between the bubble and the emulsion phase. This discrepancy is due primarily to the extensive bubble coalescence above the jet nozzle and the assumption that no gas leaks from the bubble phase. High speed movies were used to document the phenomena above a 0.4 m diameter jet in a 3-meter diameter transparent semicircular jetting fluidized bed (Yang et al., 1984b). The movies were then analyzed frame by frame to extract information on bubble frequency, bubble diameter, and jet penetration depth. The process of bubble formation is very similar to that described in Kececioglu et al. (1984), but it was much more irregular in the large 3-m bed. Because of this irregularity, it was difficult to count accurately the “bubbles” between the constrictions immediately above the jet nozzles. The initial bubble formation at the top of the jet was assumed complete when the gaseous tail with entrained solids disappeared into the gas bubbles. The largest horizontal dimension of the bubble was reported to be the bubble size; the distance between the top of the jet nozzle and the bottom of the gas bubble, the jet penetration depth. The Davidson and Harrison (1963) model assumed there was no net exchange of gas between the bubble and the emulsion phase. The validity of this assumption was later questioned by Botterill et al. (1966), Rowe and Matsuno (1971), Nguyen and Leung (1972), and Barreto et al. (1983). The predicted bubble volume, if assumed no net gas exchange, was considerably larger than the actual bubble volume experimentally observed. A model was developed to describe this phenomenon by assuming that the gas leaks out through the bubble boundary at a superficial velocity equivalent to the superficial minimum fluidization velocity. For a hemispherical bubble in a semicircular bed, the rate of change of bubble volume can be expressed as:

11-Oct-2001 JMR

Recirculating and Jetting Fluidized Beds dVB

Eq. (32)

dt

275

π DB2

= G j − U mf

2

where VB = πDB3/12 for a hemispherical bubble. Equation (32) can be reduced to show the changes of bubble diameter with respect to time in Eq. (33):

(4G

Eq. (33)

DB2 dDB j

−

2πUmf DB2

)

=

dt π

Integrating Eq (33) with the boundary condition that DB = 0 at t = 0 gives:

Eq. (34)

t =

1 2Umf

  2G + D 2π G U Gj j B j mf  ln   2πG U  j mf  2G j − DB 2π Gj Umf 

  −D  B    

The maximum bubble size, where the total gas leakage through the bubble boundary equals the total jet flow, can be obtained from either Eq. (32) or Eq. (33):

Eq. (35)

(DB )max =

2G j πU mf

The total amount of gas leakage from the bubble at a bubble size DB is: Eq. (36) F=

 2G + D 2π G U Gj  Gj j B j mf ln   2Gj − DB 2π Gj Umf 2Umf  2π Gj Umf  

  3  − D  − π DB B   12  

Equation (34) and the experimental bubble frequency, n = 1/t, were used to predict the expected bubble diameter. The predicted bubble diameters are very close to those actually observed. Theoretically, Eqs. (34) and (36) can be solved to obtain both the bubble frequency and the bubble diameter if the

11-Oct-2001 JMR

276 Fluidization, Solids Handling, and Processing total gas leakage at the moment of bubble detachment, F, is known. The only correlations available now to predict F are the empirical ones shown in Fig. 20. The bubble growth equations can be similarly derived for a circular jet in a three dimensional bed. The same experimental observation and conclusions described above for a semicircular bed are expected to hold as well.

Figure 20. Comparison between the experimental bubble flow and the total flow through the jet nozzle.

Bubble frequency from the jet was also studied using a force probe in the same bed (Ettehadieh et al., 1988). The frequency characteristics of the force probe responses and their average power spectrums were calculated. Typical average power spectrum plots of the density fluctuations measured along the bed height are shown in Fig. 21. The frequencies of the more pronounced peaks in the average power spectrum plots represent the characteristic frequencies of the periodic components of the fluctuations. Due to the presence of overlapping periodic phenomena in the lower jetting sections of the bed such as jet pulsations, bubble formation and bubble coalescence, multiple major frequencies are observed in these

11-Oct-2001 JMR


277

portions of the bed. Whereas in the middle and upper portions of the bed only a single narrow band of maximum power, or in other words, a single major frequency peak, is present. According to Broadhurst and Becker's (1976) analysis of the fluidized bed pressure fluctuations, a single maximum power peak is observed whenever a large number of random bubbles coalesce into a small number of periodic bubbles.

~ w

~18.4

3: o Q..

~~

CM--

,-~7.5CPM

~t

~~~~~-120.0CPM

..J'

1,;1~ h = 16.5

~. 0

I

I

I

I

2

4

6

8

I

CM

,

I

,

,

10 12

14

16

18

FREQUENCY.

20

HERTZ

Figure 21. Typical average power spectrum plots of the bed density fluctuations measured along the bed height with the 40.6-cm jet assembly (Set Point 9).

The axial profiles of the characteristic major frequencies measured along the bed height with the 25.4-cm and the 40.6-cm jet nozzle assemblies under different jet flow rates are shown in Figs. 22 and 23. The frequency profiles reveal that increasing the jet flow rates results in an increase in the

11-Oct-2001 JMR

278 Fluidization, Solids Handling, and Processing frequency of the oscillations measured in the lower jetting and bubble coalescing portions of the bed by in effect extending the bubble coalescence region. The figures also show that regardless of the jet flow rates and the size of the two different jet diameters, the frequency profiles all level off into a narrow frequency band at approximately 2 m above the jet inlets. All visually observed jet penetration depths fall within this jet influence zone. Bubble frequencies measured visually in the three viewing area by the frame-by-frame analysis of motion pictures are also shown in Figs. 22 and 23. Bubbling frequencies reported from the motion pictures are the frequencies of all bubbles, random or periodic, which were observed within the viewing areas. Whereas the force probe results are measured at specific bed elevations and represent the passage frequency of the most periodic single or coalesced bubbles. The slight discrepancy between the results reported by the two different sampling methods is therefore not unexpected. In the upper portions of the bed, where bubble coalescence is almost complete, the frequencies measured by the force probes are very close to the visually observed bubbling frequencies. In the lower bubble formation and coalescence portions of the bed, where bubbles form randomly and coalesce into larger bubbles, the characteristic frequencies are smaller than the overall bubbling frequencies observed visually.

~ c.. (..) >(..) z w :) c w ~ 1.1,.

Figure 22. Profiles of the characteristic major frequencies measured along the bed height with the 25.4-cm jet assembly.

9-Oct-2001 JMR


279

140

LINE TYPE SYMBOL

~ 120

~~, \

100 ~ 0.. u >u z w ~ c w ~ u.

80

60

,

, \ \, \ , I I I , I \ I +"

---

1

.\ ".

40

SET POINT

1- + ~ MOTIONPICTURES .3 o 4 () 5 .6 o 7 .8 6 9 .10 L

CJ 1- ---,

4

J---0:-4

20 0

' 0

I

I 100

I

I 200

I

I

I

300 ELEVATION.

I

I

400

I 500

I

I 600

I 700

CM

Figure 23. Profiles of the characteristic major frequencies measured along the bed height with the 40.6-cm jet assembly.

The bubbling frequencies measured in this study are much smaller than the 5– 8 Hz measured by Rowe et al. (1979) and the 20 Hz measured by Hsiung and Grace (1978). The difference may be due to the dominant effect of the much larger initial bed heights and jet nozzle sizes in this study. The size, the velocity, and the solid volume fraction of the bubbles and the solids slugs are reflected in the shape and frequency of the force probe response as shown in Fig. 24. The amplitude of the oscillations is a measure of the solid slug velocity and solid volume fraction. Inside the bubbles, the probe response is nearly flat because of the negligible solids fraction inside the bubble. The spread of the peaks and the length of the flat portions of the probe responses are measures of the size and velocity of the solid slugs and gas bubbles respectively. The average time spent by the bubbles or solids slugs in traveling from one probe location to another was measured by calculating the average time-lag between the density fluctuations as they travel along the bed height. The average bubble rise velocities calculated from the force probe measurements with 40.6–cm jet configuration are summarized in Table 1.

9-Oct-2001 JMR


w c ~ 1:J c.. ~ ~

0

2

3

4

5

6

7

8

TIME. SECOND

Figure 24. Typical local bed density fluctuations as measured by the force probe.

Table 1. Average Bubble Rise Velocities (m/s) Calculated from the Force Probe Measurements and Comparison with Results from the Movies; 40.6-cm Diameter Jet Assembly Force Probe Elevation (cm) Probe 1 Probe 2

3

321

371

321

474

2.93 (+43%)

371

523

2.69 (+31%)

321

523

4

Set Point 6

7

8

9

10

3.21 (+20%)a

4.11 (+80%)

3.21 (+46%)

4.11 (+58%)

3.66 (+39%)

5.06 (105%)

3.3 (+39%)

2.71 (+1%)

3.02 (+32%)

2.87 (+30%)

3.64 (+40%)

3.15 (+20%)

3.38 (+37%)

3.71 (56%)

2.87 (+30%)

3.00 (+15%)

2.94 (+12%)

2.92 (+19%)

3.56 (+50%)

3.05 (+16%)

2.93 (+19%)

4.06 (+70%)

2.63

2.46

2.38

3.05 (+14%)

5

2.75 (+21%)

2.79 (+27%)

Movies Results 2.05 2.68 2.28 2.20 from Overall View a ( ) refers to percent deviation from the results of movie analysis.

9-Oct-2001 JMR

2.60

Recirculating and Jetting Fluidized Beds 3.2

281

Gas Mixing Around the Jetting Region

Gas Mixing Around Single Jets. The gas exchange between the jet and the outside emulsion phase was studied by tracer gas injection and by integration of gas velocity profiles in the jet at various heights above the jet nozzle in a 28.6 cm diameter bed with a 3.5 cm jet and polyethylene beads as bed material (Yang et al., 1984a). The typical tracer concentration profiles across the bed obtained at five different elevations are shown in Figs. 25 and 26 for a nominal jet velocity of 32.6 m/s and with two different aeration flows outside the jet. The concentration profiles penetrate farther into the emulsion phase when there is less aeration flow outside the jet. The concentration profiles obtained at different elevations are also approximately similar if the local tracer concentration is normalized with the maximum tracer concentration at the axis, C/Cm, and plotted against a normalized radial distance, r/(r1/2)c, where (r1/2)c is the radial position where the tracer concentration is just half the maximum tracer concentration at the axis (see Figs. 27 and 28). Thus, we experimentally determined that, in a permanent flamelike jet in a fluidized bed, not only the velocity profiles in the jet, but also the gas concentration profiles, are similar. The gas mixing between the jetting region and the emulsion phase and the gas flow pattern around the jet were determined by solving the tracer gas conservation equation numerically along with the axial velocity profiles in the jet obtained with a pitot tube as discussed in the section, “Momentum Dissipation of a Gas-Solid Two-Phase Jet.” It is concluded that the gas mixing in a jetting fluidized bed with a permanent flamelike jet is due primarily to convection and that diffusion plays a negligible role. The resulting velocity profiles and the flow pattern inside and around the jet are shown in Figs. 29 and 30 for a jet velocity of 32.6 m/s and with two different aeration flows. The jet boundary at Vz = 8 m/s shown in Figs. 29 and 30 was calculated from Tollmien similarity. The boundary where the tracer gas concentration becomes zero, C = 0, was determined from the normalized experimental tracer gas concentration profiles shown in Figs. 27 and 28. The arrows in Figs. 29 and 30 indicate the actual axial and radial flow components, and the magnitudes are in m/s. The gas flow direction is predominantly from the emulsion phase into the jet at distances close to the jet nozzle. This flow can be from the aeration flow in the emulsion phase, as in the cases of high jet velocity or high aeration flow, or from the flow recirculated from the upper part of the jet. The

11-Oct-2001 JMR

282 Fluidization, Solids Handling, and Processing entrainment of gas into the jet occurs immediately above the jet nozzle. The extent of this region depends on both the aeration flow outside the jet and the jet velocity .Increases in aeration flow and jet velocity tend increase the height of this region. Beyond this gas entrainment region, the gas in the jet is then expelled from the jet along the jet height. The gas expelled at the lower part of the jet is recirculated back into the jet, setting up a gas recirculation pattern at the lower part of the jet. The extent of this recirculation pattern increases with increases in jet velocity and with decreasesin aeration flow outside the jet. I I I I Sampll"q ~ ~io"- I cm from Jet r«Ilzle

,

1.0

" u' u

~ 5 .5 ~

0.7

'e ~u

0.6

~ u .. ~ --

O.S

'0 ~

0..

~ ~

o

~

9.3

.23.6 A 33.8 a ~.

"'0 \

0.8

.1.7

45.5

3. ~.

7 & t«I. 9 Fbw

JetVeklClly = 32.6mls \,

\' \ \,

~ 4\,

\\ \ r \ \ , 4 \, \' ~ , \, \, , \ 'p , , , , \ 4 I', , ~. i ,I' \ \D \ ,

,\ \ \

~~ I

L

0.3

0.2

~.3

FkJW

\.1' \ ~. 9 FkIw

7 FkJW

.\ 0.1

~, ~, \ \,

\

, , -I~

0 0

2

.

6

8

10

12

14

Back Will 16

Distance tram Front Plate. cm

Figure 25. Experimental concentration profiles at different distances from the jet nozzle (Run GJ18).

11-Oct-2001 JMR


283

Figure 26. Experimental Concentration Profiles at Different Distances from the Jet Nozzle (Run GJ35).

9-Oct-2001 JMR


SamplIng Port location cm from Jet t«»zzle .1.7 o 9.3

1.0

.23.6 6 33.8

~ 0.9

a

45. S

t«». 3. t«».7.& ~. 9 Fk>w Jet Veklclty = 32.6 m/s 0.1

0.7

E

0.6

u u

0.5

0.4

0.3

0.2

0.

0 0

1.0

z.o

(r/rl/2'C

Figure 27. Normalized concentration profiles for Run GJ18.

9-Oct-2001 JMR

3.0


285

Figure 28. Normalized concentration profiles for Run GJ35.

9-Oct-2001 JMR


287

Figure 30. Calculated axial and radial gas velocity profiles for Run GJ35.

9-Oct-2001 JMR

288 Fluidization, Solids Handling, and Processing The axial velocity profiles, calculated on the basis of Tollmien similarity and experimental measurement in Yang and Keairns (1980) were integrated across the jet cross-section at different elevations to obtain the total jet flow across the respective jet cross-sections. The total jet flows at different jet cross-sections are compared with the original jet nozzle flow, as shown in Fig. 31. Up to about 50% of the original jet flow can be entrained from the emulsion phase at the lower part of the jet close to the jet nozzle. This distance can extend up to about 4 times the nozzle diameter. The gas is then expelled from the jet along the jet height.

Figure 31. Calculated gas entrainment into a jet in a fluidized bed based on Tollmien’s Similarity.

9-Oct-2001 JMR


289

Gas Mixing Around Concentric Jets. Gas mixing phenomena around a concentric jet were investigated by Yang et al. (1988) in a large semicircular cold flow model, 3 meters in diameter and 10 meters high, with a triple concentric jet nozzle assembly of 25 cm in diameter. In this discussion, the outermost annulus in the triple concentric jet assembly is called the shroud, the innermost jet is called the feed line, and the middle annulus, the air tube. Solids were fed into the bed only through the feed line. The jet velocities employed range from 15 to 60 m/s and solid feeding rates in the concentric jets, from 0 to 3000 kg/hr. The experiments were carried out by injecting carbon dioxide as the tracer gas selectively into various flow streams and taking gas samples across different cross-sections of the bed. The bed material was -8+70 mesh crushed acrylic particles with a harmonic mean diameter of 1406 µm and a particle density of 1200 kg/m 3. A dividing gas streamline was observed experimentally which prevents the gas mixing between the jetting region and the emulsion phase until at higher bed heights. This dividing gas streamline corresponds roughly to the boundary of downflowing solids close to the walls. Similar observation was also noted in a smaller bed as discussed in the section “Gas Mixing

Around Single Jets.” Typical contour profiles of equal tracer gas concentrations are shown in Figs. 32–34. In Fig. 32, the zero percent tracer concentration boundary represents the extent of the dispersion of the gas that was injected into the fluidized bed via the air tube. Increases in solids feeding rate improve the gas mixing. Similarly, the contour plots that appear in Fig. 33 show that in the initial jet expansion region, no contribution from the tracer gas that was injected through the conical grid could be detected. This means that there is no entrainment of gas from the conical section by the concentric jet flows. Partial entrainment of gas from the conical region into the core of the reactor takes place only at upper portions of the jetting region, i.e., at locations where bubbles form and coalesce. Figure 34 shows that the region of influence for the narrow shroud flow jet is limited to its immediate surrounding and that there is substantial lateral mixing of the shroud and the emulsion phase gases. Several observations were made based on this study. Regardless of the incoming jet flow rate, the gases that are injected through the concentric jets essentially remain in the core of the reactor and do not fully mix with the gas in the dense solid downflow region of the bed. Similarly, the gas injected through the conical grid sections is not entrained by the incoming

24-Sep-2001 JMR


/

400 I ,

350

I

/"

"

; /

,1 \1

, ... I

.

-/

/

./ ~

°II

i

1.-I"

I

IH 300 I\ Ii

I

I \ I \ ~ IiI ! I I :E 250 IIi = I I

f II

I I I 10

Z8rD I

,

I I I I

I I r

,

1

1 ,~\111

f

\\ \ '

1\

\

5q

\

\1

I

I

I

I

\

m, \

\ \1

\

\

,\

\

I

I

I I

\

,

~ ~ =

r 1

1

1

II

~\

I

n

\~

~150

~

,

~111 01111

,III 1 --\ \ 501111I

100

II 180 \ \ I-\~\

111111 1II1I I 11111I 111\11

\ \\ UI \\ 80

\

\ \\\\\\\ I' I'l

I ,

I

l\

-\

-.1'/111'1

~ ~

i

M =

1

1

1111\1

11"1

,

\\

"

\

\

1\

~

;

=

'

,1

200 ~

'E

~

-a

.;!

=

9

::

\ I \

~

/

/'

,1"1'\\111 ."11\\\\\\\1

1

, \

II'

\

r.;"i1111

50

\11111

1

Test

"

CO2 Air . Feed Feed Long

0

III

" of Grid

.d

Polnl4 Tube = 60.96 mlsec , = 15.54 mlsec ..JIB = O kg/hr

0-

Region

LCO2 Jel

Figure 32. Tracer gas contour profiles—tracer injected via air tube.

9-Oct-2001 JMR


JetNozzles Figure 33. Tracer gas contoUr profiles-tracer

injected via conical grid.

291

292

Figure

Fluidization,

Solids Handling,

34. Tracer gas contour profiles-tracer

and Processing

injected via shroud.


293

jets. Partial entrainment and mixing of these gases occur at locations where bubble formation and bubble coalescence take place. On the contrary, the mixing among the concentric jets occurs quite fast, usually completed within the jet penetration length. Typical axial profiles of the tracer gas concentration measured at three different radial locations above the feed line, the air tube, and the shroud are compared in Fig. 35. In this test, the carbon dioxide tracer gas was injected in the air tube flow. As expected, there is considerable mixing between the jets as they enter the fluid bed. After an initial jet mixing zone, roughly corresponding to the jet penetration depth, the three jet flows become fully mixed. Above this inner jet mixing zone, the concentration profiles continue to fall as they mix with the gas from the emulsion phase. This is the region where the bubble formation and most of the bubble coalescence take place. Above this region, the jets and the emulsion phase gases in the vicinity of the reactor core become well mixed and the concentration profiles level off. At these bed locations regular bubble passage was visually observed during bed operation.

1.0

1:= O ... m :: 1:= 0) 1.) c o u 1/1 m (:)

~ 1.) m ~

0.6

0.4

0.2

0 0

40

80

120

160

Distance

200 Above

240

280

320

360

400

Jet Inlet. cm

Figure 35. Typical axial tracer concentration profiles—tracer gas injected via air tube at air tube gas velocity of 31 m/s.

11-Oct-2001 JMR

294 Fluidization, Solids Handling, and Processing The dividing streamline, which prevents the gas mixing between the jets and the emulsion phase until at higher bed heights shown in Figs. 32 through 34 as the zero contour lines, corresponds roughly to the boundary of downflowing solids close to the wall. The measurement of solids flow patterns with a force probe was discussed in details by Yang et al. (1986) and will be presented in the next section. Axial tracer concentration profiles measured with a higher air tube gas velocity are shown in Fig. 36. The mixing zone for the three concentric jets is larger due to the higher jet velocity. The effect of feed line solids loading on gas mixing of concentric jets can also be seen in Fig. 37. It clearly indicates that the mixing rate among the concentric jets increases with increasing feed line solid loading. Changes in bed height, however, have little effect on gas mixing.

Figure 36. Typical axial tracer concentration profiles—tracer gas injected via air tube at air tube gas velocity of 61 m/s.

11-Oct-2001 JMR


295

Figure 37. Effect of solids loading on gas mixing among concentric jets—tracer injected via air tube at air tube gas velocity of 31 m/s.

3.3

Solids Circulation in Jetting Fluidized Beds

The solids circulation pattern and solids circulation rate are important hydrodynamic characteristics of an operating jetting fluidized bed. They dictate directly the solids mixing and the heat and mass transfer between different regions of the bed. In many applications, the performance of fluidized beds is frequently controlled by the hydrodynamics phenomena occurring in the beds. Applications such as the fluidized bed combustion and gasification of fossil fuels are the cases in point. In those applications, the rates of fuel devolatilization and fines combustion are of the same order of magnitude as the mixing phenomena in a fluidized bed. The mixing and contacting of the gases and solids very often are the controlling factors in the reactor performance. This is especially true in large commercial fluidized beds where only a limited number of discrete feed points for reactants is allowed due to economic considerations. Unfortunately, solids mixing in a fluidized bed has not been studied extensively, especially in large commercial fluidized beds, because of experimental difficulties.

11-Oct-2001 JMR

296 Fluidization, Solids Handling, and Processing Solids Circulation Pattern. Yang et al. (1986) have shown that, based on the traversing force probe responses, three separate axial solids flow patterns can be identified. In the central core of the bed, the solid flow direction is all upward, induced primarily by the action of the jets and the rising bubbles. In the outer regions, close to the vessel walls, the solid flow is all downward. A transition zone, in which the solids move alternately upward and downward, depending on the approach and departure of the large bubbles, was detected in between these two regions. Solids mixing and circulation rate were studied in a fluidized bed, 3 m diameter and 10 m in height, by pulse injection of tracer particles with characteristics similar to those inside the bed but with sizes larger than those in the bed. By taking solids samples continuously at different bed locations and analyzing by sieving, the rate of particle mixing and circulation can be calculated. Experiments were conducted in two separate bed configurations. One employed a 0.25 m jet nozzle assembly, a deep-bed configuration with a bed height of 5.5 m; a 0.41 m jet nozzle assembly was used in the second configuration. Crushed acrylic particles, -6 mesh and 1,400 µm in weight-mean average size and 1,100 kg/m3 in density, were used as the bed material. The bed was first operated at the preselected conditions at a steady state; then about 455 kg of the coarse crushed-acrylic particles, similar to that used as the bed material but of sizes larger than 6-mesh, were injected into the bed as fast as possible to serve as the tracer particles. Solids samples were then continuously collected from five different sampling locations at 30-second intervals for the first 18 minutes and at 60-second intervals thereafter. The samples were then sieved and analyzed for coarse tracer particle concentration. Typical tracer particle concentration profiles vs. time at each sampling location are presented in Figs. 38–42 for set point 3. Typically it took about 160 to 200 seconds to inject a pulse of about 455 kg coarse tracer particles into the bed pneumatically from the coaxial solid feed tube. It can be clearly seen from Figs. 38 to 42 that the tracer particle concentration increases from essentially zero to a final equilibrium value, depending on the location of the sampling port. The steady state was usually reached within about 5 minutes. There is considerable scatter in the data in some cases. This is to be expected because the tracer concentration to be detected is small, on the order of 4%, and absolute uniformity of mixing inside a heterogeneous fluidized bed is difficult to obtain.

11-Oct-2001 JMR


297

Figure 38. Experimental solids mixing data and model predictions—0.254 m jet, Set Point 3, Sampling Port A.

0

JXI

~

m Cumul.tie

Time.

13X)

l~

laX!

s~onds

Figure 39. Experimental solids mixing data and model predictions—0.254 m jet, Set Point 3, Sampling Port B.

11-Oct-2001 JMR


Figure 40. Experimental solids mixing data and model predictions—0.254 m jet, Set Point 3, Sampling Port C.

Figure 41. Experimental solids mixing data and model predictions—0.254 m jet, Set Point 3, Sampling Port D.

11-Oct-2001 JMR


0

m

~

~ 1200 Cumulalve Time. ss:onds

1m

299

1!XI

Figure 42. Experimental solids mixing data and model predictions—0.254 m jet, Set Point 3, Sampling Port E.

The solids circulation patterns were investigated with a force probe developed in-house. Typical force probe responses are presented in Figs. 43 and 44 for a probe located at 0.13 m from the jet nozzle and with different penetrations into the bed for an air tube velocity of 45.7 m/s. Since the force probe is directional, the upward solids movement will produce a positive response from the probe and vice versa, the magnitude of the response being an indication of the magnitude of solids circulation rate. The number of major peaks per unit time is closely related to the actual bubble frequency in the bed. The force probe data allow the identification of three major solids flow regions in the 3-m model, as shown in Fig. 45. At the central portion of the bed, the solids flow is induced upward primarily by jetting action at the lower bed height and by large bubbles at the higher bed height. At the outer region next to the vessel wall, the solids flow is all downward. The region has a thickness of approximately 0.25 m. Between these two regions the solids flow is alternatively upward and downward, depending on the approach and departure of large bubbles. No stagnant region was evident

11-Oct-2001 JMR

300 Fluidization, Solids Handling, and Processing anywhere in the bed. Figure 45 was constructed on the basis of the force probe signal. At the central portion of the bed where the solids flow was upward, the probe gave a positive signal while the signal in the outer region was all negative because the solids flow was predominantly downward. The radial position where the first negative signal was observed was taken as the boundary between the central region and the intermediate region. The probe penetration where the probe signal became all negative was taken as the transition between the intermediate region and the outer region.

.05 F

.10:

E c o ; ::. ~ ~ 11. QJ -8 ct

.15

.20

.25

.30

j Ii

J.lltlJj

.lIJ

~ .111.1 11

~.!l1li

j 1.11 J

1l1.1Lh

,1!1-JA)l

h

A, i

2

IliiJ

JIA .lild

J J

llllt.ll...LI..lll..l

J 111-

" III

L A ..Jy iA.J 1! A !~lllili:-tl

Ai&1 .i ~ lAr llllA

A4

~

.--1 Chart

Figure 43. Force probe responses for probe penetration from 0.05 m to 0.30 m— 0.13 m from jet nozzle elevation, 46 m/s jet velocity, no solid feed.

11-Oct-2001 JMR


301

Figure 44. Force probe responses for probe penetration from 0.36 m to 0.56 m— 0.13 m from jet nozzle elevation, 46 m/s jet velocity, no solid feed.

In addition to the three solids circulation regions readily identifiable, the approximate jet penetration depth and bubble size can also be obtained from Fig. 45. The jetting region can be taken to be the maximum average value of jet penetration depth. From the jet boundary at the end of the jetting region, an initial bubble diameter can be estimated. This value can be taken to be the minimum value of the initial bubble diameter. The diameter of a fully developed bubble can be obtained from the bubble boundary in the developed-bubble region, as shown in Fig. 45. The central region is thus divided further into three separate regions axially: the jetting region, the bubble-developing region, and the developed–bubble region. Bubbles were observed to coalesce in the bubble–developing region during analysis of the motion pictures taken through the transparent front plate.

11-Oct-2001 JMR

302 Fluidization, Solids Handling, and Processing BedSurface

I 4.00

Vessel 'Wall

DevelopedBu~le

175

Region

3." 125

r I

100

, 2.7S

I

I

Bu~le Devel~ing Region

.. "11

( SolIds

I

Move

I /

A"ernately

2.~

Upward

and

~nward In I ThisRegionl, ~ 2. 25 -a 8" 1!

I ( Solids

Move

I

Predominantly Downward In This R!9lonl

I

I

I

2.00 -Precbminantlv, Upward In This R89lonl

I I

1.75

I I

l.~ ~

I

I

I I

1.25 Jet

Conical Grid

1.00 ! Jetting Region

vUnda~1

I

I

75 I o

50! 25 Jet

Nozzle O

I .25

.~

.75

1.00

1.25

I. ~

1.75

Figure 45. Three major solids flow regions in PI.8. 3-m meters model derived from force Distance from front probe signal. Three major solids flow regions in 3-m model derived from force signal.

11-Oct-2001 JMR


303

Solids Circulation Rate. The results of the force probe measurement indicate that the solid circulation rates increase with increasing jet flow rates. A simple mechanistic model was developed to correlate the solids circulation data. The model assumes that the solids circulation inside the bed is induced primarily by the bubble motion. The solids circulation pattern inside the bed can be divided into two major regions radially. In the center of the bed, the particle movement is predominantly upward and the movement is induced by the bubbles disengaged from the central jet. This region has a radius similar to the radius of the average bubble size. In the outer region, the particles move primarily downward. In the meantime, the particles in both regions exchange with each other across the neighboring boundary at a constant rate of Wz g/cm2-s. This mechanistic model is shown schematically in Fig. 46. Material balance in a differential element dz as shown in Fig. 46 gives:

Figure 46. Mechanistic model for solids mixing.

11-Oct-2001 JMR

304 Fluidization, Solids Handling, and Processing In the bubble street region Eq. (37)

K

∂ X ′J ∂z

(

)

+ π Ri2 1 − ε mf ρ s

∂ X ′J ∂t

+ 2π Ri Wz ( X ′J − X J ) = 0

In the annular region Eq. (38)

K

∂ XJ ∂z

(

)(

)

+ π Ro2 − Ri2 1 − ε mf ρ s

∂ XJ ∂t

+ 2π Ri Wz ( XJ − X ′J ) = 0

where Eq.(39)

K = nVB f w (1 − ε w )ρ s

The data do not show any clear dependence on the axial position, z. The axial dependence is thus assumed to be negligible. Equations (37) and (38) are reduced from partial differential equations to ordinary differential equations. If we consider only the annular region, Eq. (38) reduces to dX J

Eq. (40)

dt

+

π

(

2π RiW z

Ro2

)(

)

− Ri2 1 − ε mf ρ s

( X J − X ′J ) = 0

Since both XJ and X´J are independent of z, the relationship between XJ and X´J can be approximated by material balance of the coarse particles injected into the bed to serve as the tracer. Solving for X´J , we have

Eq. (41)

X ′J =

Wt

π

Ri2 H

  2   1 − ε mf R −  o  − 1   1 −ε ρ s  Ri  i  

(1 − ε ) mf

 X  J 

Substituting Eq. (41) into Eq. (40), and after some mathematical manipulating, we get Eq. (42)

11-Oct-2001 JMR

dX J dt

+ PX J − Q = 0


305

where Eq. (43)

P=

Eq. (44)

Q=

(R π

2 RiW z 2 o

)(

)

− Ri2 1 − ε mf ρ s

(

+

2Wz

Ri (1 − ε i )ρ s

2Wz Wi

Ri Ro2

−

Ri2

)H (1 − ε )(1 − ε )ρ mf

i

2 s

Equation (42) can be readily integrated, however, there are two limiting cases to consider. Case I. Instantaneous Injection of Tracer Particles. If it is assumed that the tracer particles are injected instantaneously, Wt = Wto = a constant, Eq. (42) can be integrated with the boundary condition that XJ = 0 at t = t0 to give

Eq. (45)

XJ X Jo

= 1 − exp[− P(t − t o )]

Case II. Uniform Injection of Tracer Particles. Since the large amount of tracer particles usually required more than 75 seconds to inject, the other limiting case would be to assume that the injection rate was uniform over the injection period, or

Eq. (46)

W o Wt =  t  t  w

 ⋅t  

Again, Eq. (42) can be integrated with the boundary condition that XJ = 0 at t = 0 to give Eq. (47)

XJ X Jo

=

1 twP

{Pt − [1 − exp(− Pt )]}

The equilibrium tracer concentration in the bed after complete mixing can be expressed as

Eq. (48)

X oJ =

[

Wt

(

)(

π Hρ s Ri2 (1 − ε i ) + Ro2 − Ri2 1 − ε mf

)]

11-Oct-2001 JMR

306 Fluidization, Solids Handling, and Processing The voidage inside the bubble street, ε i, can be calculated as follows Eq. (49)

(

)

ε i = ε mf + f B 1 − ε mf = ε mf +

nVB πRi2U A

(1 − ε ) mf

where f B is the volumetric fraction of bubbles occupying the bubble street region at any instant; it can be evaluated from the following equation Eq. (50)

fB =

nVB π Ri2U A

If the bubble frequency, bubble diameter, and bubble velocity are known, the solids mixing rate can be calculated. The mechanistic model developed in the last section is applied to the data collected experimentally. Bubble diameter and bubble velocity calculations were based on the empirical equations obtained from frame-byframe analysis of high-speed motion pictures taken under the respective operating conditions (Yang et al., 1984c). The equations used are: For 0.254 m jet nozzle assembly Eq. (51)

Bubble diameter

DB = 12 .36 ⋅ G 0. 155

Eq. (52)

Bubble velocity

U A = 0. 711 gDB

For 0.406 m jet nozzle assembly Eq. (53)

Bubble diameter

DB = 0. 0195 ⋅ G 0. 620

Eq. (54)

Bubble velocity

U A = 0 .35 gD

In correlating the data, the solid exchange rate between the two regions, Wz, was assumed to be constant. The tracer concentration data were analyzed statistically and the solids circulation rates are reported in Table 2. The positive fluxes indicate that the net solids flow is from bubble

11-Oct-2001 JMR


307

street to annular region. The net exchange fluxes do not seem to depend on the jet velocity and the bed configuration. The solids circulation rate depends on the jet velocity, however, because higher jet velocities generate larger bubbles. The circumferential area surrounding the bubble street will then be larger and thus the solids circulation rate will be larger. The solids circulation rates derived on the basis of this model range from 47,500 to 73,400 kg/h, as shown in Table 2. Comparison of the calculated and the experimentally observed tracer concentration profiles is good as shown in Figs. 38 through 42 for set points 3 employing the 0.254 m jet nozzle assembly. The solids mixing study by injection of tracer particles indicated that the axial mixing of solids in the bubble street is apparently very fast. Radial mixing flux depends primarily on the bubble size, bubble velocity, and bubble frequency, which in turn depend on the size of the jet nozzle employed and the operating jet velocity.

Table 2. Statistical Analysis of Solids Mixing Data

Test No.

t0 s

0.254 m Jet Nozzle Assembly Set Pt. 1 8.3 Set Pt. 2 Set Pt. 3 Set Pt. 4

44.5 10.2 - 10.2

Wz kg/m2 s

2Ri Wz (H - J) kg/s

1.0

13.2

1.3 1.3 1.0

16.2 17.7 13.4

1.0 0.84

14.7 20.4

0.406 m Jet Nozzle Assembly Set Pt. 1 Set Pt. 2

- 1.2 - 8.6

11-Oct-2001 JMR

308 Fluidization, Solids Handling, and Processing Solid Entrainment Rate into Gas and Gas-Solid Two-Phase Jets. A mathematical model for solid entrainment into a permanent flamelike jet in a fluidized bed was proposed by Yang and Keairns (1982). The model was supplemented by particle velocity data obtained by following movies frame by frame in a motion analyzer. The experiments were performed at three nominal jet velocities (35, 48, and 63 m/s) and with solid loadings ranging from 0 to 2.75. The particle entrainment velocity into the jet was found to increase with increases in distance from the jet nozzle, to increase with increases in jet velocity, and to decrease with increases in solid loading in the gas-solid, two-phase jet. It is well known that jets in a fluidized bed induce high solids mixing. In one extreme, jets can be permanent and flamelike, similar to jets in spouted beds. The solids circulation in this case is created by solids entrainment into the jet along the jet height. Studies of solids circulation in spouted beds have been reviewed by Mathur and Epstein (1974). Data on solids entrainment into a permanent flamelike jet immersed in a fluidized bed, however, are meager. In another extreme, jets can be a series of rapidly coalescing bubbles, called bubbling jets. Solids mixing in this case is induced essentially by the solids-carrying capacity of the bubble wake and by the bubble frequency. Another kind of jet encountered in operating fluidized beds are those created by pneumatic transport of solid particles into the fluidized beds. Here we call them gas-solid, two-phase jets because the incoming jet streams have already entrained solid particles at different loadings. The momentum of these solid particles is not negligible, as already shown by Yang and Kearins (1980). With high-speed movies such particles can usually be seen to penetrate right through the roof of coalescing bubbles in the bubbling jet regime. Regular and high-speed movies were taken of the tracer particle movement around the jets at different velocities and different solid loadings. The tracer particles used are red plastic pellets of similar size and density to the bed material. The movies were then analyzed frame by frame using a motion analyzer to record the particle trajectories and the particle velocities. Typical particle trajectories observed in the movies are shown in Fig. 47 for a jet velocity of 62.5 m/s and a solid loading of 1.52. The time elapsed between dots shown in Fig. 47 was typically 5 movie frames, while the movie speed was 24 frames/s. The colored tracer particles were followed in the vicinity of the jet until they disappeared into the jet, as

11-Oct-2001 JMR


309

indicated by an arrow into the jet in Fig. 47. Sometimes the tracer particles disappeared into the bed before they reached the jet. In those cases, no arrow is shown in Fig. 47. If the points where the particles disappeared into the jets were connected, a jet expansion angle could be readily determined. The angle determined from Fig. 47 is approximately 15°. This gives a jet half-angle of 7.5°. Other sets of data also give a comparable jet half-angle.

E u

...

I t

~ Z "Q; -. E o .. ... ..

II

~

~ Q

,v \t~

.Ii Time Elaps81BetweenDots ...Is 5 Ml7ile Frames

I .r:

...~ I' -16

,

, -lZ

,

, -8

Olstance

,

, -4

/Jet ...,

, O

From Center

4

NoZZle , 8

I

, lZ

I

,

I

16

01 Jet, cm

Figure 47. Particle trajectories around a jet in a fluidized bed.

11-Oct-2001 JMR

310 Fluidization, Solids Handling, and Processing The average particle velocity in each trajectory was determined by dividing the distance traveled by the particle with the total time elapsed. The particle velocity into the jet increased with the distance from the jet nozzle, decreased with increases in solid loading, and increased with increases in jet velocity. High-speed movies indicated that the entrained particles tended to bounce back to the jet boundary more readily under high solid loading conditions. This may explain why the entrainment rate decreases with increases in solid loading in a two-phase jet. A ready analogy is the relative difficulty in merging into a rush-hour traffic as compared to merging into a light traffic. The dependence of particle velocity on the distance from the jet nozzle is presented in Fig. 48 for a jet velocity of 62.5 m/s and a solid loading of 1.52. The regressional analysis of this dependence at different operating conditions is summarized in Table 3.

Figure 48. Experimental solid velocity into the jet.

11-Oct-2001 JMR


311

Table 3. Summary of Experimental Solids Velocity into the Jet at Different Operating Conditions ______________________________________________________________________________ Run No.

Nominal

Solid Loading

Regressional Relationship

Correlation

Jet Vel. Wt. Solid/Wt. Gas w.r.t. Dist. from Jet Nozzle Coefficient (m/s) (Vz = cm/s; Z = cm) ______________________________________________________________________________ GSF-1

1.21

0

Vz = 0.0788 Z + 4.707*

0.8158

GSF-3

1.21

0.40

Vz = 0.0688 Z + 5.265

0.4433

GSF-4

1.21

0.92

Vz = 0.0842 Z + 2.979

0.8959

GSF-5

1.21

1.52

Vz = 0.0765 Z + 2.692

0.9085

GSF-22

1.20

0

Vz = 0.1811 Z + 3.119

0.8705

GSF-23

1.20

0.51

Vz = 0.1212 Z + 3.597

0.8439

GSF-24

1.20

1.21

Vz = 0.1839 Z + 2.228

0.8953

GSF-25

1.20

1.99

Vz = 0.1115 Z + 2.000

0.9593

GSF-44

1.19

0

——

——

GSF-45

1.19

0.71

——

——

GSF-46

1.19

1.67

——

——

GSF-47 1.19 2.75 Vz = 0.0967 Z + 2.576 0.8750 ______________________________________________________________________________ * Vz = solid particle velocity, cm/s; Z = distance from jet nozzle, cm

Model for Solids Entrainment into a Permanent Flamelike Jet. A simple model for solids entrainment into a permanent flamelike jet is described here. The jet is assumed to expand at an angle 2θ as shown in Fig. 49, where θ is commonly known as the jet half-angle. Although the existence of a jet half-angle for the jet in a fluidized bed is not universally accepted, employment of this concept considerably simplifies the development of the model. The concept may also be applicable to a bubbling jet (Anagbo, 1980). Material balance of solid particles in a differential element inside the jet gives Eq. (55)

Wj + dWj − Wj = Vjz (2π r ⋅ dz )(1 − ε z )ρ s

or Eq. (56)

dWj = Vjz ρs (1 − ε z )2π r ⋅dz

11-Oct-2001 JMR

312 Fluidization, Solids Handling, and Processing where Wj = solids circulation rate, Vjz = horizontal component of the particle velocity into the jet at z and ε z = voidage of emulsion phase at z.

Figure 49. Schematic of jet for solids entrainment model development.

Through geometric consideration, we find that Eq. (57)

r = z ⋅ tanθ +

do 2

where d o is the diameter of the jet nozzle.

11-Oct-2001 JMR


313

Substituting Eq. (57) into Eq. (56), we have Eq. (58)

 d dW j = 2πρs (1 − ε z )V jz  z ⋅ tan θ + o  2

  dz 

The overall entrainment rate into the jet or the rate of solid circulation induced by the jet is then

Eq. (59)

Lj  d  Wj = 2π ρs (1 − ε z ) ∫V jz  z ⋅ tan θ + o  dz 0  2 

where Lj is the jet penetration depth and ε z is assumed to be independent of jet height. We found experimentally that the particle velocity into the jet, Vjz, is linearly dependent on the jet height, as expressed below: Eq. (60)

Vjz = C1 z + C2

where C1 and C2 are two empirical constants. Some of those constants at different jetting conditions were reported in Table 3. Substituting Eq. (60) into Eq. (59) and integrating, we have Eq. (61)  C tanθ 3 1  C1d o   C d W j = 2πρ s (1 − ε z ) 1 L j +  + C2 tanθ L2j + 2 o L j  2 2 2   3 

The correlation for the jet penetration depth, Lj , was developed earlier as shown in Eq. (23). The solids entrainment rate into a jet in a fluidized bed can be calculated from Eqs. (61) and (23) if the empirical constants C1 and C2 and the jet half-angle θ are known. The jet half-angle θ can be taken to be 10° as suggested by Anagbo (1980), a value very close to 7.5° obtained from solid particle trajectories reported here. The real jet half-angle will be larger than 7.5° because of the truncation of the jet by the front plate of the semicircular bed. The particle velocity can usually be approximated as shown in Eq. (27). When the jet velocity is low or the bed particles are relatively fine or of wide size distribution, the jet tends to be a bubbling jet. A separate model

11-Oct-2001 JMR

314 Fluidization, Solids Handling, and Processing for solids circulation is required. The present model can be used as a first approximation. The model as formulated in this section cannot be used to predict a priori the solids entrainment rate into the jet because of the two empirical constants in Eq. (61). Lefroy and Davidson (1969) have developed a theoretical model based on a particle collision mechanism for entrainment of solid particles into a jet. The resulting equation for particle entrainment velocity is

Eq. (62)

Ve

=

Vj

π 2e(1 + e )d p 16r

assuming the d p /2r

-pJ

0

Q5

III 1.0

1.5

2.0

I 2.5

3D

Normalized Diameter

Figure 40. Self-preserving size distributions of different materials. (From Sastry, K. V. S., Int. J. Min. Processing, 2:187–203, 1975, with kind permission from Elsevier Science-NL, Sara Burgerhartstraat 25, 1055 KV Amsterdam, The Netherlands.)

11-Oct-2001 JMR

Coating and Granulation

415

0.9 0.8

0.7 ~ ~

0.6

. ~

-a ~

0.5

0.4

0.3 0.2 0.1 0

2

4

6

8

10

Figure 41. Evolution of variance and average size based on a cut off kernel, with w = (uv)2/ (u + v)3. (From Adetayo and Ennis, 1996.)

11-Oct-2001 JMR

416 Fluidization, Solids Handling, and Processing An Example of Modeling Coalescence in Granulation Processes. Growth by coalescence in granulation processes may be modeled by the population balance, as mentioned. It is necessary, however, to determine both the mechanism and kernels which describe growth. For fine powders within the non-inertial regime of growth where granule deformation can be neglected, all collisions result in successful coalescence provided that binder is present. Therefore, coalescence occurs via a random, or size independent kernel which is only a function of liquid loading y, or Eq. (42)

β (u , v) = k = k * fn (y)

The dependence of growth on liquid loading represented by fn(y) strongly depends on wetting properties. For random growth, it may be shown that the average granule size is given by Eq. (43)

a = ao e k t

where a0 is the initial nuclei size, and kt is the extent of granulation. Random growth will occur until the inertial regime of granulation is reached, at which point an additional relationship exists to determine the largest granule size possible, or

Eq. (44)

St =

16 ρ U o amax 9µ

= St*

Combining the previous two relationships allows us to estimate the extent of granulation as a function of the governing group of the Stokes number, or

Eq. (45)

 *  µ  ( k t )max = 6 ln St fn ( y ) ∝ ln  St o   ρ U o ao 

Hence, the extent of non-inertial growth depends logarithmically on binder viscosity and the inverse of agitation velocity (Adetayo et al., 1995). Maximum granule size depends linearly on these variables. Also, the extent of growth has been observed to depend linearly on liquid loading y and,

11-Oct-2001 JMR


417

therefore, maximum granule size, exponentially on liquid loading. Figure 42 illustrates the normalization of extent of growth for the drum granulation of limestone and fertilizers, using the data of Kapur (1972).

(a)

Figure 42. Extent of granulation as a function of fractional saturation. (From Adetayo, et al., Powder Technol., 82:37–49, 1995, with kind permission from Elsevier Science S.A., P.O. Box 564, 1001 Lausanne, Switzerland.) (a) Extent of granulation, k 1t1, as a function of fractional saturation and (b) extent of granulation normalized for differences an Stvo as a function of fractional saturation. (Inset) Comparison to data of Kapur (1972): ¡, AS; ¨, MAP; ∆, DAP; n, limestone (Kapur, 1972).

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Unwanted Aggregation in Fluidized Beds

Industrial fluidized beds are operated with particles which usually contain impurities or are themselves a mixture of many components. Upon heating, some of these components soften, melt or react with each other, generating stickiness. In fluid bed reactors, the fluidizing gas as well as the solid particles can react chemically producing new components and, during this process, particle cohesion and stickiness can occur. In all these cases, the fluidization behavior of the bed changes dramatically as temperature is increased and unwanted agglomerates form. Such characteristics as the minimum fluidization and bubbling velocity, the bubble size and bubble frequency, all undergo a significant change and, in the limit, total defluidization of the bed can take place. Like in granulation, stickiness generated by surface softening or due to the presence of a sticky chemical species results in agglomerate formation and growth. However, unlike granulation, this growth is usually unintended and it proceeds in the noninertial regime where, as long as stickiness is present, coalescence takes place and can, in some cases, become uncontrollable. It was shown (Siegell, 1984; Compo et al., 1984; Tardos et al., 1985a,b; and Compo et al., 1987) that cohesiveness and subsequent agglomeration and sintering of fluidized particles occurs if fluid beds are operated at temperatures at or above the so-called minimum sintering temperature (point), Ts, of the particles. It was also shown that for pure materials fluidized in inert gases, the sintering temperature is an intrinsic property of the solid particle surface which can be estimated using constant heating rate dilatometry (Tardos et al., 1984). This temperature is usually a fraction of the melting point of the solid material and depends on structure and chemical composition; Table 9 gives a few examples for materials used by Compo et al. (1987). As can be seen from the table, the ratio of the sintering point to the melting temperature varies a great deal and can take values from approximately 0.5 to 0.95. These results show that “the rule of thumb” employed in industry for the ratio Ts/TM to be approximately 0.8 is correct for some compounds but not for others. Using the “dilatometer” technique, a small sample of powder (about 1– 2 grams) is heated at constant rate in the apparatus depicted schematically in Fig. 43. Dilatation of the sample is measured by a linear voltage transducer (LVDT); contraction of the sample indicates particle-particle surface flattening and defines the minimum softening point or sintering temperature, Ts. In

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Table 9. Dependency of Minimum Sintering Temperature on the Electronic Structure of Crystalline Material

Formula

T m(° C)

T s(°C)

T s/T m (K/K)

Sodium Chloride

NaCl

802

400

0.63

Sodium Bromide

NaBr

755

420

0.67

Calcium Chloride

CaCl2

772

440

0.68

Calcium Fluoride

CaFl2

1330

368

0.39

Magnesium Oxide

MgO

2800

1109

0.45

Polyethylene Granules

125

92

0.92

Polypropylene Beads

162

141

0.95

Polyethylene Beads

135

118

0.96

polyethylene Spheres

135

127

Material Ionic Compounds

Covalent Network Compounds

Aluminum Nitride

Al2N 2

Zeolite Cracking Catalyst

0.98 *

2150

1108

0.57

1260

900

0.77

Metals Copper Shot (dp = 1.015 mm)

Cu

1083

899

0.86

Copper Shot (dp = 0.718 mm)

Cu

1083

844

0.83

Copper Shot (dp = 0.056 mm)

Cu

1083

750

0.75

Nickel Powder

Ni

1452

434*

0.41

Mo

2620

814

*

0.40

Boron Powder

B

2330

1759*

0.78

Silicon Powder

Si

1420

1335*

0.95

Molybdenum Powder Covalent Compounds

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420 Fluidization, Solids Handling, and Processing the same instrument, dilatation or contraction at constant temperature (isothermal dilatometry) can also be performed; this procedure yields particle surface flattening kinetics or interparticle contact deformation rates. These experiments are usually performed in an inert atmosphere such as nitrogen or helium. The results from these types of measurements can be used to predict sinter-neck growth and hence allow prediction of the rate of strengthening of formed agglomerates at constant temperature.

WEIGHTS

LVDT

SAMPLE CUP UPPER PISTON

PARTICULATE SAMPLE oSAMPLE HOLDER

LOWER PISTON

Figure 43. Schematic diagram of a dilatometer. (From Tardos and Pfeffer, Powder Technology, 85:29–35, 1995, with kind permission from Elsevier Science S.A., P.O. Box 564, 1001 Lausanne, Switzerland.)

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A more accurate but also much more difficult and time consuming experiment to determine the sintering temperature and other fluidization properties of a given material at elevated temperatures is carried out in the three-inch diameter, temperature controlled fluidized bed depicted in Fig. 44, using 1–2 kg of fluidizable powder. Defluidization is accompanied by a large rapid change in pressure drop through the bed which can easily be detected. It was found previously that the formation of large agglomerates and bed defluidization occur at temperatures somewhat higher than the minimum sintering temperature determined in the dilatometer experiments described above. This is clearly due to the fact that in the dilatometer particles touch at fixed points while in the fluid bed some particle movement takes place and particle contacts are necessarily intermittent.

Figure 44. Schematic diagram of the high temperature fluidized bed reactor/ granulator. (From Tardos and Pfeffer, Powder Technology, 85:29–35, 1995, with kind permission from Elsevier Science S.A., P.O. Box 564, 1001 Lausanne, Switzerland.)

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422 Fluidization, Solids Handling, and Processing It was also shown both theoretically and experimentally (Tardos et al., 1985a,b) that there is a strong correlation between the excess temperature above the minimum sintering point, T-Ts, and the excess gas velocity above minimum fluidization conditions (measured below the sintering point) U-U mf, required to maintain fluidization. A general correlation was developed between the excess temperature and the excess gas velocity, which takes the form

U -Um f Eq. (46)

Umf

D  T -Ts  =K B d p  Ts 

n

where DB and dp = 2a are the bed and particle diameters, and K and n are material dependent coefficients. These take different values for amorphous and crystalline materials and are given in Table 10.

Table 10. Values of Coefficients K and n

Material

K

n

Amorphous Coal Ash, Glass, Polymer

0.1

2

Crystalline Salts, Ores

0.0025

0.4–0.5

There is a significant difference in the form of Eq. (46) for amorphous and crystalline materials (Compo et al., 1987 and 1990); for the first group, a slight increase in temperature above minimum sintering results in a significant increase in excess gas velocity as can be seen in both, the value of the

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constant K as well as the value of the exponent (n = 2, in this case). For crystalline materials, on the other hand, the correlation is weaker and the exponent is significantly smaller (n = 0.4–0.5). The determination of sintering temperatures was extended to more complex materials such as coal ashes (Ladesma et al., 1987) and beneficiated ilmenite ores (Compo et al., 1987). Although these experiments were also performed in an inert atmosphere, different components within these materials reacted and sintered at different temperatures. It was again found that sintering and softening detected in the dilatometer predicted very closely the agglomeration and defluidization observed in the experimental fluid bed in accordance with the results given in Eq. (46). During more recent work (Tardos and Pfeffer, 1995) several systems were studied where fluidizing gas actually reacted with fluidized particles and either a solid or a gaseous product was formed. The influence of product formation on fluidization was studied in both an isothermal fluidized bed and in the dilatometer. The following reactions were studied: reduction of calcium sulfate (phosphogypsum) to calcium sulfide (CaS) in hydrogen, oxidation of coke and magnesium in air and production of aluminum nitride (AlN) from a solid precursor containing aluminum and carbon by carbothermal nitridation (Nicolaescu et al., 1994). In one case, the product was a gas (CO 2) while in others it was a solid (calcium sulfide, aluminum nitride and magnesium oxide). One of the products was a fine, extremely cohesive material (CaS) while the others were stable, free flowing materials which exhibit no cohesiveness (MgO and AlN). It was shown that in the case of the gaseous product, the bed was easily fluidizable as long as the temperature did not exceed the sintering point of the solid reactant. When the product was a noncohesive, solid powder, steady fluidization could be maintained until total conversion was achieved by keeping the temperature below the sintering point of both the reactant and the product. Production of the cohesive powder, calcium sulfide, was only possible in a fluid bed diluted with heavy inert particles (sand) which acted as fluidization media and promoted the removal of the product from the surface of the reactant powder. In the case of the aluminum nitride production, even though carbon (coke) is not readily fluidizable at temperatures above 1200– 1300°C, the formation of a fluidizable product (AlN) ensured good fluidization to temperatures as high as 1500–1600°C.

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424 Fluidization, Solids Handling, and Processing ACKNOWLEDGMENT The authors would like to acknowledge that the section dealing with population balance modeling of granulation processes is an abbreviated version of material prepared by Dr. J. D. Litster, University of Queensland, for a joint short course given by Dr. B. J. Ennis and Dr. J. D. Litster.

NOTATIONS a

Effective radius or characteristic length of a particle

m

A

Area of indentor in indentation test

m2

B

Birth rate of particles of a given size due to coalescence, Eq. (5) or nucleation, Eq. (40)

# particles/kg·s

B’

Dimensionless concentration, Eq. (3)

-

c

Crack length

m

Ca

Capillary number = Uo µ/γ

-

Co

Laplace-Young pressure deficiency

-

Ddef

Coefficient of deformation of a granule

Cp,g

Specific heat capacity of gas

D

Death rate of particles by coalescence, Eqs. (5) and (40)

J/kg·°C

DB

Bed or bubble diameter

DG

Dispersion coefficient for growth rate (growth diffusivity)

# particles/kg·s m kg2/s

d

Indentor diameter

m

dp

Particle diameter

m

dp,o

Initial droplet diameter

m

e

Coefficient of restitution of a solid surface

-

E

Expectation (expected or average value) or Young’s Modulus

Pa

E’

Elutriation rate, Eq. (10)

kg/s

f(m)

Mass distribution of particles

kg-1

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Coating and Granulation fsur

Fraction of surviving granules

F

Total force

N

Fvis

Viscous force from Eq. (18)

N

Fcap

Capillary force from Eq. (19)

N

Fcoat

Feed rate of coating material (solute only)

kg/s

g

Acceleration due to gravity

m/s2

G

Growth rate of particles due to coating

kg/s

Gc

Critical strain energy release rate

J/m2

h

Height of wetted powder in a Washburn Test or half separation distance between two particles held in contact by a pendular liquid bridge

m

Minimum separation distance between two particles due to surface asperities

m

hb

Height of a fluidized bed

m

he

Equilibrium height of rise in Eq. (15)

m

H

Hardness (from indentation test)

Pa

ha

425

k, k1, Constants in Eqs. (7, 8 and 46) K, K’ kg

Thermal conductivity of gas

W/m·°C

Kc

Fracture toughness of material

Pa·m1/2

l

Wear displacement

m

m

Mass

kg

MWg

Molecular weight of gas

kg/kmol

MW l

Molecular weight of liquid

kg/kmol

n

Coefficient in Eq. (46)

nk

Mass distribution of feed or product particles

n(v,t)

Number frequency distribution

N

Total number of granules or number of particle circulations

kg-1

# particles

pl*

Vapor pressure of liquid

Pa

ptot

Total pressure of system

Pa

P

Applied load in indentation test

N

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426 Fluidization, Solids Handling, and Processing Qk

Rate of feed or product withdrawal

# particles/s

rp

Characteristic length of process zone

m

R

Radius of capillary

m

S

Saturation or spray rate

-

St, St* Stokes number and critical Stokes number, Eq. (20)

-

Sf

Shape Factor in Eqs. (29) and (31)

-

t

Time

s

Tbulk

Bulk gas temperature

°C

Ts

Surface temperature of liquid droplet or minimum sintering temperature of a solid

°C

TM

Melting point of a solid

°C

∆T

= (Tbulk - Ts)

°C

tv

Time for complete evaporation of a liquid droplet

s

tcirc

Time elapsed between successive passages of a particle through the spray zone

s

u

Particle size measure (volume) for coalescing particles m3

U

Superficial gas velocity

m/s

UB

Rise velocity of a bubble

m/s

Ui

Impact velocity

m/s

Umf

Minimum fluidizing velocity

m/s

Uo

Relative velocity of two approaching particles

m/s

v

Particle size measure (volume) for coalescing particles m3

V

Volumetric wear or breakage rate or volumetric flow of feed

Var

Variance

w*

Critical average granule size capable of rebound from a collision

Xtotal

Total mass of coating received by a particle in a batch coating operation

xc

Critical specimen size, Eq. (35)

xk

Amount of mass received by a particle in its kth passage through the spray zone

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m3/s

kg

kg

Coating and Granulation y

Liquid loading

ybulk

Mass fraction of solvent in the bulk

-

yd

Mass fraction of solvent in the liquid droplet

-

ys

Mass fraction of solvent in gas at the liquid surface

-

Y

Yield number, Eq. (28)

-

427

Greek Letters α

Thermal diffusivity

m2/s

β

Constant in Eq. (34)

-

β()

Coalescence kernel, Eq. (40)

-

Γ

Shear rate

γ

Surface tension or interfacial energy

δ

Dimensionless bubble spacing

-

δc

Effective increase in crack length

m

ε

Dimensionless separation of two particles = 2h/a or powder void fraction, Eq. (26)

θ

Contact angle between liquid and powder compact

λ

Latent heat of vaporization

µ

Viscosity

kg/m·s

ρ

Density

kg/m3

σy

True yield stress of material

Pa

σf

Applied fracture stress

Pa

τ

Coating run time

τo

Yield strength of wet granule

φ

Filling angle defined in Fig. 22

ϕ

Probability distribution function

s-1 N/m

J/kg

s N/m2

Subscripts ave

Average value

breakage Breakage c

Crystal

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428 Fluidization, Solids Handling, and Processing circ

Circulation

d

Droplet

e

Equilibrium, exit

g

Gas or granule

i

Inlet

I-C

Inertial to coating transition

l

Liquid

max

Maximum value

N-I

Non-inertial to inertial transition

p

Particle

tot

Total

wear

Wear

yy

Denotes stress value in material

Superscripts C

Coating

fb

Fluid bed

I

Inertial

lv

Liquid-vapor

N

Non-inertial

sv

Solid vapor

sl

Solid-liquid

*

Critical value

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Ritala, M., Jungersen, O., Holm, P., Schaefer, T., and Kristensen, H. G., “A Comparison Between Binders in the Wet Phase of Granulation in a High Shear Mixer,” Drug Development and Industrial Pharmacy, 12:(11–13), 1685–1700, Marcel Dekker, Inc. (1986) Ritala, M., Holm, P., Schaefer, T., and Kristensen, H. G., “Influence of Liquid Bonding Strength on Power Consumption during Granulation in a High Shear Mixer,” Drug Development and Industrial Pharmacy, 14(8):1041– 1060 (1988) Robinson, T., and Waldie, B., “Particle Cycle Times in a Spouted Bed of Polydisperse Particles,” Can. J. Chem. Eng., 56:632–635 (1978) Robinson, T. and Waldie, B., “Dependency of Growth on Granule Size in a Spouted Bed Granulator,” Trans. I. Chem. E., 57:121–127 (1979) Rumpf, H., Particle Technology, Chapman and Hall, London (1975) Sastry K. V. S., “Similarity Size Distribution of Agglomerates During Their Growth by Coalescence in Granulation or Green Pelletization,” Int. J. Min. Processing, 2:187–203 (1975) Sastry K. V. S., and Fuersteau, D. W., “Mechanisms of Agglomerate Growth in Green Pelletization,” Powder. Tech., 7:97–105 (1973) Schaefer, T., “Equipment for Wet Granulation,” Acta Pharm. Soc., 25:205–228 (1988) Schaefer, T., Holm, P., and Kristensen, H. G., “Wet Granulation in a Laboratory Scale High Shear Mixer,” Pharm. Ind., 52(9):1147 (1990) Schaefer, T., Holm, P., and Kristensen, H. G., “Melt Granulation in a Laboratory Scale High Shear Mixer,” Drug Development and Industrial Pharmacy, 16(8):1249–1277 (1990) Siegell, J. H., “High Temperature Defluidization,” Powder Technol., 38:13–22 (1984) Smith, P. G., and Nienow, A. W., “On Atomizing a Liquid into a Gas Fluidised Bed,” Chem. Eng. Sci., 37:950–954 (1982) Tardos, G. I., Mazzone, D., and Pfeffer, R., “Measurement of Surface Viscosities using a Dilatometer,” Can. J. of Chem. Eng., 62:884–887 (1984) Tardos, G. I., Mazzone, D., and Pfeffer, R., “Destabilization of Fluidized Beds Due to Agglomeration, Part I: Theoretical Analysis,” Can. J. of Chem. Eng., 63:377–383 (1985a) Tardos, G. I., Mazzone, D., and Pfeffer, R., “Destabilization of Fluidized Beds Due to Agglomeration, Part II: Experimental Verification,” Can. J. of Chem. Eng., 63:384–389 (1985b) Tardos, G. I., Pfeffer, R., and Ennis, B. J., “A Microlevel-Based Study of Granulation,” Proc. of the Powder and Bulk Solids Conf., pp. 221–230, Rosemont, Illinois (1991)

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434 Fluidization, Solids Handling, and Processing Tardos, G. I., Garcia, J., Ahart, R., and Ratuiste, F., “How and Why Some Flow Aids Promote Bulk Powder Flow,” AIChE Symp. Series, 296(89):18–34 (1993) Tardos, G. I., “Measurement of Powder and Binder Characteristics which Determine Agglomerate Growth during Granulation,” Proc. of the 25th Israel Conf. on Mech. Eng., Haifa, Israel (1994) Tardos, G. I., “A Review of High Shear Mixer Granulation,” Proc. of the First Israel Conf. for Conveying and Handling of Particulate Solids, Herzliya, Israel (1995) Tardos, G. I., and Khan, M. I., “Study of Granulation in a Constant Shear Granular Flow Couette Device,” Paper presented at the Annual AIChE Meeting, Miami Beach, Florida (1995) Tardos, G .I., and Pfeffer, R., “Chemical Reaction Induced Agglomeration and Defluidization of Fluidized Beds, Powder Technology, 85:29–35 (1995) Tardos, G. I., Klan, M. I., and Mort, P. R., “Critical Parameters and Limiting Conditions in Binder Granulation of Fine Powders,” Powder Tech., 94:245– 258 (1997) Tsutsumi, A., Nakmoto, S., Mineo,T., and Yoshida, K., “A Novel Fluidized-Bed Coating of Fine Particles by Rapid Expansion of Supercritical Fluid Solutions,” Proc. 1st Int. Particle Technol. Forum, pp. 452–455, Denver, CO (1994) Waldie, B., “Growth Mechanism and the Dependence of Granule Size on Drop Size in Fluidized-Bed Granulation,” Chem. Eng. Sci., 46:2781–2785 (1991) Waldie, B., and Wilkinson, D., “Measurement of Particle Movement in a Spouted Bed Using a New Microprocessor Based Technique,” Can. J. Chem. Eng., 64:944–949 (1986) Wan, L. S. C., and Lai, W. F., “Factors Affecting Drug Release from Drug-Coated Granules Prepared by Fluidized-Bed Coating,” Int. J. Pharmaceutics, 72:163–174 (1991) Weiss, P. J., and Meisen, A., “Laboratory Studies on Sulphur-Coating Urea by the Spouted Bed Process,” Can. J. Chem. Eng., 61:440–447 (1983) Wurster, D., U.S. Patent 2,799,241 (1957) Wurster, D., U.S. Patent 3,253,944 (1966) Xu, M., “A Theoretical and Experimental Investigation of Spouted Beds with Draft Tubes,” Ph.D. Dissertation, West Virginia University, Morgantown, WV (1993) Yuregir, K. R., Ghadiri, M., and Clift, R., “Impact Attrition of Sodium Chloride Crystals,” Chem. Eng. Sci., 42:843–853 (1987)

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7 Attrition in Fluidized Beds and Pneumatic Conveying Lines Joachim Werther and Jens Reppenhagen

1.0

INTRODUCTION

Attrition of particulate materials occurs wherever solids are handled and processed. In contrast to the term comminution, which describes the intentional particle degradation, the term attrition condenses all phenomena of unwanted particle degradation which may lead to a lot of different problems. The present chapter focuses on two particular process types where attrition is of special relevance, namely fluidized beds and pneumatic conveying lines. The problems caused by attrition can be divided into two broad categories. On the one hand, there is the generation of fines. In the case of fluidized bed catalytic reactors, this will lead to a loss of valuable catalyst material. Moreover, attrition may cause dust problems like explosion hazards or additional burden on the filtration systems. On the other hand, attrition causes changes in physical properties of the material such as particle size distribution or surface area. This can result in a reduction of product quality or in difficulties with operation of the plant.

435

436 Fluidization, Solids Handling, and Processing Both, the mechanism and the extent of particle degradation depend not only on the process type but also on properties of the solid material, and to a large extent on the process conditions. Clift (1996) has stated that attrition is a triple-level problem, i.e., one is dealing with phenomena on three different length and time scales: the processing equipment, the individual particles, and the sub-particle phenomenon such as fracture which leads to the formation of fines. The appearance of attrition can, therefore, differ very much between the various applications. For that reason, the following section deals with the various modes of attrition and the factors affecting them. In order to evaluate the extent of attrition and its impact on the particle size distribution, there is a need of a qualitative and quantitative characterization. This, however, is not as simple as it may seem at first. There are many different properties, parameters and effects that manifest themselves and could be measured. In addition, as will be shown, the choice of the assessment procedure is strongly connected with the definition of attrition which, on its part, depends on the degradation mechanism that is considered to be relevant to the process. Hence there are a lot of procedures and indices to characterize the process of particle attrition. Section 3 deals with those which are relevant to fluidized beds and pneumatic conveying lines. Unfortunately, the basic physical mechanisms that control the attrition process are still poorly understood. As a consequence, particular test methods are used to evaluate the degradation tendency of the materials or to predict the rate of attrition for a given process. There are a lot of procedures using widely different devices and operations. Some of them observe the degradation of only one individual particle, whereas others treat a considerable amount of material. The particles are subjected to stress systems which range from well-defined ones like impact or compression, to those which are similar to the more or less randomized stresses occurring in natural processes. Section 4 attempts to summarize the huge variety of attrition tests in a systematic way. Test facilities are also used to investigate specific attrition phenomena of individual processes. Up to now only a few investigations have been carried out in full-scale equipment. Most results were obtained from very special devices which makes it difficult to compare the results of various research groups and to draw general conclusions. Despite of these difficulties, the specific attrition phenomena in fluidized beds and pneumatic conveying lines will be summarized in the Secs. 5 and 6, respectively. It will

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Attrition in Fluidized Beds and Pneumatic Conveying Lines 437 be shown that the fluidized bed systems can be divided into several regions which are characterized by completely different attrition characteristics. These regions will be discussed separately with regard to the respective influence of the operating and design parameters. Some models and correlations to predict or at least describe the extent of attrition will also be given. On the other hand, the number of papers relating to attrition in pneumatic conveying lines is so small that the results can be outlined in a qualitative manner only. At the end of both sections some general rules will be presented to reduce the attrition in a given case.

2.0

FACTORS AFFECTING ATTRITION

Before discussing the individual factors affecting attrition in detail, the distinction has to be made between two completely different modes of attrition, namely fragmentation and abrasion. The two modes and their different effects on the particle size distribution are sketched in Fig. 1. Abrasion produces a lot of very fine material removed from the surface of the initial particles that retain their identity and become only slightly smaller. Abrasion leads to a bimodal size distribution. On the other hand, fragmentation destroys the particles and produces a number of particles which are all distinctly smaller than the original ones. The resulting size distribution becomes broader and has a smaller mean diameter than the original one. In industrial processes the two modes usually do not occur separately. They are rather combined in varying proportions. There may be, for example, an initial breakage of mother particles followed by surface abrasion of the resulting fragments. In this case a small number of intermediately sized particles and many fine particles will be produced. In general, the extent of abrasion in relation to fragmentation depends on the various factors discussed below. However, there have been several rules suggested to identify the dominant mode in a given case. They all use the change in size distribution as criterion. For instance, Dessalces et al. (1994) assessed the degradation behavior of various industrial FCC catalysts by the change in the ratio ∆dp3,90 /∆ dp3,10, i.e., the ratio of the 90% and 10% values of the cumulative mass density distribution of the particle sizes. They got high values of this ratio (>100) if abrasion was dominant and smaller values (

0.75

=

E

2 E

rlR

(r/R)

+

0.1'11

b.47 E

3.62(r/R)

+ 0.191

Figure 27. Radial Voidage Distribution for FCC Catalyst in a 90 mm i.d. column, 2.25 m above gas distributor (curves computed, points experimental) (Tung, Li and Kwauk, 1988).

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Bubbleless Fluidization 6.0

537

SHALLOW FLUID BEDS

Another mode of bubbleless gas/solid contacting originated from the identification of a shallow region above the fluid bed distributor, before bubbles take shape. By far the major portion of gas/solid transfer takes place in this narrow region, and the contribution of the overlying solids is often only nominal. A critical factor in designing shallow beds is an understanding of the functioning of the fluid bed distributor to insure even gas distribution and minimal pressure drop. 6.1

Dynamics for the Distributor Zone

Figure 28 (Yan, Yao and Liu, 1982) presents the results of turbulence measurement above a distributor plate by means of a hot-wire anemometer, indicating that turbulence intensity increases sharply near the distributor, especially for the perforated plates and the cap plate, and then diminishes beyond a certain distance (40 mm in the present case) to some nearly constant low values. Figure 29 (Qin and Liu, 1982) shows the behavior of individual particles above the distributor recorded by video camera of small clusters of particles, coated with a fluorescent material and spot-illuminated by a pulse of ultra violet light from an optical fiber. The sequential images, of which Fig. 29 just represents exposures after stated time intervals, were reconstructed to form the track of motion of the particle cluster shown in Fig. 30. Neither this track nor visual observation of the shallow bed while fluidized, reveal any vestige of bubbles. Instead, the particles are thrown up by the high velocity jets issuing from the distributor orifices to several times their static bed height. 6.2

Activated Solids Shallow Fluid Bed Heat Exchanger

Figure 31 (Liu, Liu, Li and Kwauk, 1986) shows a cylindrically shaped shallow fluid bed tubular heat exchanger. Solid particles are fluidized with a small stream of activating gas Ga, so as to insure maximal heat transfer between the particles and the exchanger tube wall. The waste gas Gw, from which heat is to be extracted, passes through the solid

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Bubbleless Fluidization

Figure 29. Motion of luminescent particles in shallow fluidized bed. (Qin, and Liu, 1982.)

539

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150

100

" ~

k

~

50

~

\~

u

~

~

Figure 30. Track of particle motion in shallow fluidized bed. (Qin, and Liu, 1982.)

Figure 31. Activated Solids Shallow fluid-bed heat exchanger. (Liu, Liu, Li, and Kwauk, 1986.)

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541

particles already suspended by the activating gas, thus heating the solid particles. In merely percolating through the suspended particles, the waste gas experiences minimal pressure drop. The waste gas does not pass through the gas distributor, and therefore, cannot possibly clog the distributor holes for the carbonaceous materials the hot gas may contain. Figure 32 shows the variation of pressure of both gas streams, Ga and Gw, with their flow ratio. Obviously the pressure of Gw is much less than that of Ga, even at a flow ratio N = Gw/Ga of as high as 10. 6.3

Cocurrent Multistage Shallow Fluid Bed

For slow reactions, the shallow fluid beds have been organized into a cocurrent multistage fluid bed (MSFB) reactor as shown in Fig. 33 (Yan, Yao, Wang, Liu and Kwauk, 1983). In this reactor, solids are carried up by the flowing gas stream, and once they reach the top, they are collected through a funnel and recirculated to the bottom by means of a pneumatically controlled downcomer. Rather unexpected flow behavior was discovered while the shallow fluid beds were connected in series for cocurrent solids flow with the gas, as shown in Fig. 34. When a shallow dense bed is present at any stage, there is a positive increase of bed height to any increase in solids flow rate, at any given gas velocity. However, if the solids rate is reduced to some critical value, the dense bed slumps to practically nil height. If the solids rate is restored gradually, the dense bed does not build up until another (upper) critical value is reached. Then the dense bed reappears all of a sudden. Between these two critical solids flow rates, there exists a hysteretic region in which there are two stable bed heights corresponding to any solids rate: a dense bed and a dilute bed. This hysteresis suggests two different mechanisms of solids transport from a lower stage to an upper stage: elutriation in the dense phase, and “jetting” in the dilute phase. Above the upper critical solids flow rate, only dense bed exists, and below the lower critical rate, only dilute bed exists, both these extreme regions being monotonic in the relation between bed height and solids flow rate.

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IU ~ a CI..

Figure 32. Variation of pressure of activating gas and waste gas with their flow ratio. (Liu, Liu, Li, and Kwauk, 1986.)

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543

Figure 33. The cocurrent multistage (co-MSFB) fluid-bed reactor. (Yan, Yao, Wang, Liu, and Kwauk, 1983.)

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545

The Co-MSFB as a Chemical Reactor

The viability of the co-MSFB as a chemical reactor for slow reactions has been examined for the gaseous reduction of iron ores. The experimental setup is shown in Fig. 35 (Yan, Yao, Wang, Liu and Kwauk, 1983), the heart of which is a 70-mm i.d. reactor, either fitted with a 13stage co-MSFB internal, or empty. Comparative reduction tests for equal inventories of iron ore revealed that the reduction time required for any stated degree of metallization could be reduced considerably by the use of the co-MSFB internals, as shown in the set of curves of Fig. 36. Whether the advantage of the staging accrues from better G/S contacting, or from the suppression of back flow of the gas stream from stage to stage, remains yet to be corroborated from mass-transfer studies between the solids and the flowing gas.

Figure 35. Experimental co-MSFB reactor for iron oxide reduction. (Yan, Yao, Wang, Liu, and Kwauk, 1983.)

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Figure 36. Comparison of rate of reduction of iron oxide with and without coMSFB (Yan, Yao, Wang, Liu, and Kwauk, 1983).

7.0

FLUIDIZATION WITH NO NET FLUID FLOW

Certain reactions between solids go through intermediate steps in the gas phase. For instance, the segregating roasting of oxidic copper ores with addition of salt and coal, and the reduction of metal oxides with solid carbon. These reactions, known as pseudo-solid-solid reactions, may sometimes be carried out advantageously in the fluidized state. Solid fluidization would enhance transport of the gaseous intermediates, and would facilitate the addition and withdrawal of a circulating thermal carrier in the case of an endothermic reaction. Inasmuch as no gaseous reactant needs to be introduced, it would be logical to seek means of achieving fluidization, even on a periodic basis, with no net fluid flow.

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547

Levitation of Discrete Particles

Levitation is a stable condition in which a particle responds to the oscillating fluid in such a way that the influence of finite buoyancy, or gravity, forces is completely neutralized so that the particle oscillates about a fixed position (Houghton, 1963, 1964, 1966, 1968; Krantz, Carley and Al-taweel, 1973; Tunstall and Houghton, 1968; Van Oeveren and Houghton, 1971). Figure 37 (Liu, 1983) shows the levitation of solid particles in air under oscillation caused by a sonic generator located at the bottom of the column.

(a)

(b)

Figure 37. Levitation of solid particles in vertically oscillating air. (a) Dilute suspension of solids. (b) Solids concentration at node. (Liu, Keling, 1983.)

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548 Fluidization, Solids Handling, and Processing A simplified analysis (Deng and Kwauk, 1990) of solids levitation led to the following ordinary differential equation involving three parameters, u x, du x /dx, and d 2u x /d 2x, to be determined from experiments:

(

)

du p π 3 π d pρ p = − C D d 2p ρ f u p −u f u p −u f + 6 dt 8

π + d 3p ρ f 6

 du f du p  du p  u x + − +  dt  dr  dt

[(

)

(

)d

](

+ u f − u p u x +u p u f −u p −v u f −u p

) du

x

dx

−

ux  π 3  − d p ρ p −ρ f g dx 2  6 2

(

)

in which u p is the velocity of the particles, u f is the velocity of the piston creating oscillation, and u x is the function of x in the variable-separated form of the fluid velocity u(t,x). Experiments for verifying the above equation were carried out in a Ushaped plexiglas tube, shown in Fig. 38a (Deng and Kwauk, 1990), one leg of which is filled with particles for oscillation tests and voids itself into a large-diameter tank at the top which both serves as a hydraulic seal for the system and also dissipates the kinetic energy of the oscillating fluid. Fluid oscillation was generated by a mechanical device, shown in Fig. 38b. Ionexchange resin particles, 0.29 to 0.44 mm in diameter, were tested in water, and 0.44 mm glass beads in glycerol-water mixtures. Figure 39 shows the visual appearance of resin particles in oscillating water at different frequencies. Figure 39a shows that at low frequency the particles merely jump up and then fall back to the supporting wire net. Figure 39b shows that as frequency increases, the particles may remain in the liquid for some time before falling down. Figure 39c shows that at some still higher frequency, the particles remain levitated, executing vertically oscillating motion. At even higher frequencies the particles will rise against gravity (not shown in Fig. 39).

11-Oct-2001 JMR

(\') ri

~ ~ ... ...

O O

QJ ~ ~

... ... .2j .2j

TI vv

uu

"0 "0 cc

.-O .-O S-;: S-;: Q. Q.

~ ~ uu

>. u

~ ~

o -I/) o -I/) ..~E~..c ..~E~..c

eo..eo..-

O O

...z:.., u"O~"O Q;e~ ~ ; .. u

e e ~ 0

=r

c o ~ s "> .!! ..c Q)

~

~'

0.Q) 0. .=

0 v,

-1

~

QJ ~

... QJ

.5

~

~

549

11-Oct-2001 JMR

Figure 38. Liquid\solid levitation apparatus. (a) U-Shaped fluid oscillation tube. (b) Mechanical oscillation generator. (Deng and Kwauk, 1990.)


/ GI .D B

0&&

-0

GI "' p,~ tU ~ .c ."' >< I GI ;:J-a.

(b) (a)

..,.


11-Oct-2001 JMR

(a)

(b)

(c)

Figure 39. Ion-Exchange resin in vertically oscillating water (particle diameter 0.435 mm, amplitude 2 cm, asymmetric factor 1.2). (a) Frequency f = 3.08 Hz; (b) f = 4.82 Hz; (c) f = 5.40 Hz. (Deng, and Kwauk, 1990.)


551

The frequency of fluid oscillation at which levitation takes place is plotted in Fig. 40 against the corresponding amplitude A of oscillation, the asymmetry factor k o or ratio of the duration of the downstroke to that of the upstroke, and the resin particle diameter d p. From these experimental data, the three parameters of the equation were correlated to the particle diameter: u x = 11.1d p + 0.536

du x dx d 2u x dx

2

= −39 .5d p + 2.20

= − 1.15 × 106 d p + 7. 25 × 104

Curves computed from the equation using these empirically correlated parameters are also shown in Fig. 40, indicating acceptable agreement. Trajectories of a single resin particle computed from the equation are shown in Fig. 41, indicating that at f = 4.5 Hz the particle sinks, while at f = 7.5 Hz rising levitation takes place. 7.2

Semi-Fluidization through Oscillatory Flow

Both Fig. 37 and Fig. 39 show solid particles highly dispersed as a dilute phase in the oscillating fluid, either gas or liquid, without evidence of any bubbles. When a packed bed of solid particles was subjected to the action of an oscillating liquid, however, only the upstroke portion of the periodic fluid motion was capable of dispersing the solid particles against the action of gravity, while during downstroke they fell back onto the distributor plate. The dynamics of such a mode of semi-fluidization is similar to that of jigging in ore dressing, which is a common operation in coal or ore dressing, though little used in the design of chemical reactors. Thus, at least on a periodic basis, jigging yields the same advantages as other modes of fluidization with no net fluid flow.

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11-Oct-2001 JMR Figure 40. Comparison of computation from the equation on p. 548 with experimental data for resin-water system. (Deng, and Kwauk, 1990.)


553

E v. ~

Figure 41. Trajectories of single resin particle in water computed from the equation on p. 548 (particle diameter 0.435 mm, amplitude 2 cm, asymmetric factor 1.2). (Deng, and Kwauk, 1990.)

7.3

Application to Pseudo Solid-Solid Reactions

The pseudo solid-solid reduction of oxidic iron ores with solid carbon is known to proceed through the gaseous intermediates, CO and CO2, according to the following mechanism: → CO + FeO ↑ CO2 + C ←

→ CO2 + Fe → 2CO ↓ → vent

At low temperatures the overall rate is controlled by the chemical reaction CO2 + C → 2 CO, but at sufficiently high temperatures the controlling steps would be taken over by mass transfer, which could be materially accelerated by oscillatory gas flow. This pseudo solid-solid reduction of oxidic iron ores with solid carbon could be carried out in a conceptualized jigged reducer shown in Fig. 42 (Kwauk, 1979, 1981, 1983). Ore and coke are fed as a mixture at a single

11-Oct-2001 JMR


Figure 42. Conceptual design of a jigged iron ore reducer (Kwauk, 1979).

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555

point to the middle section of a three-section reducer. A considerable amount of CO is generated between reaction of C and FeO, which fluidizes the two top sections by continuous flow. Division into the two fluidized sections helps the establishment of a solids concentration gradient (dilute in the top section, and dense in the lower section) which tends to suppress longitudinal mixing. Only the lowest of the three-section reducer is semifluidized through jigging, in which only the last traces of CO are evolved. All the CO resulting from the pseudo solid-solid reaction is conducted, together with entrained char, from the top fluidized section through a constriction, in which the high-velocity gas flow prevents backflow, to a transport combustor, where the CO is burned to CO2 with preheated air, along with as much of the char as is called for by heat balance to maintain the endothermic FeO-C reaction. The heated recycled char is separated from the off gas at the top of this transport combustor in a hot cyclone and is returned as a thermal carrier to the lower part of the lowest jigged section, while the hot flue gas from the transport combustor is used to preheat the incoming air in a recuperator. In the top fluidized section, where the first portion of FeO reduction, that is, prereduction, takes place, the lighter coke particles are elutriated from the heavier ore particles. The lower fluidized section is sized for adequate FeO-C mixing in order to maximize contacting of the reacting solids. Then, in the lowest section, where vestigial amounts of CO are evolved, solids are semi-fluidized through jigging to effect an additional stage of coke upflow in order to denude the powdered metallic iron product of its accompanying char. Jigging is produced by a gas oscillation generator, which could be either mechanical or electromagnetic, separated from the bottom of the jigged section first by a solids knockout hopper and then by a thermal break, which is essentially a packed-bed heat regenerator. To protect further the gas oscillation generator from hot gases, a cooler is interposed between it and the thermal break. The process proposed in Fig. 42 uses powdered iron ore directly without prior treatment such as sintering or pelletization as called for in the case of the blast furnace, and powdered coke or coal without prior gasification. Jigging through gas oscillation replaces expensive gas recycling which calls for first cooling and then reheating, often accompanied by unavoidable carbon deposition through the Bouduord reaction.

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PARTICLES WHICH QUALIFY FOR BUBBLELESS OPERATION

Whether or not a solid powder qualifies for bubbleless gas/solid contacting depends greatly on its fluidizing characteristics, which may be determined in a bed collapsing apparatus which automatically tracks the subsiding bed surface of the collapsing bed by means of an optical-fiber probe (Qin and Liu, 1985). Modeling of the collapsing process (Tung and Kwauk, 1982; Yang, Tung and Kwauk, 1985; Tung, Yang, Xia, Zheng, Yang and Kwauk, 1989) resulted in a dimensionless number to quantify the fluidizing characteristics of powders. Certain binary and ternary solid mixtures have been shown to demonstrate synergistic actions in improving the fluidization characteristics of the component particles, implying the significance of particle size design. 8.1

Powder Characterization

Broadly speaking, for G/S systems, three modes of particle-fluid contacting may be recognized to take place simultaneously as shown in Fig. 43: bubbles containing sparsely disseminated particles, emulsion of densely suspended particles, and defluidized (transient as well as persistent) particles not fully suspended hydrodynamically by the flowing gas. For all intents and purposes, it is desirable to suppress bubbles and to prevent defluidization. When the gas, which is fluidizing a powder exhibiting all these modes of contacting, is turned off abruptly, the fluidized bed will collapse and subside in three consecutive stages, as shown in Fig. 44 (Yang, Tung and Kwauk, 1985): (i) A rapid initial stage for bubble escape (ii) An intermediate stage of hindered sedimentation with constant velocity of the dense emulsion of mobile particles, and (iii) A final decelerating stage of solids consolidation for the incompletely suspended particles. Thus, the random spatial distribution of the three modes of particlefluid contacting is transformed into the ordered, temporal sequence of the three stages of the sedigraph.

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557

Figure 45 shows the instrument used for automatic tracking of the surface of the collapsing bed, inclusive of a computer for data acquisition and on-line analysis. The fluidized bed, 5 cm in diameter and 120 cm high, is provided with a high pressure-drop gas distributor to insure uniform gas flow. Below the distributor is a specially designed knife valve operated by a solenoid for quick gas shufoff. An optical-fiber probe, consisting of two separate sets of projector and receiver fibers, respectively, for illumination and for reception of reflected light, is used for rapid tracking and recording of the subsiding bed surface. Figure 46 shows a typical bed collapsing curve traced by the instrument described above. The entire determination, from gas shutoff to printout from the computer, rarely exceeds 3 to 4 minutes for normal solids. And since all measurements are taken by the instrument, they are not subject to personal error of observation. The size of solids test sample is of the order of a kilogram or less.

Figure 43. The three modes of particle-fluid contacting in G\S fluidization.

11-Oct-2001 JMR


1=0 (i)

z.

(i)

---(ii)

z.-

stagel:bubble

escape

(iii)

.. v 01 't: " "'

(iv) stage2:hindered

sedimentation

"0 .8 (v) critical

z.

I I I

z.

---~-D-

O t. )

I.

point

stage3:solidsconsolidation

~~~~e-

~ii)

I I

t.

:-t-~1

'V

time

t-~

--

Figure 44. Modelling the three-stage bed collapsing process. (Yang, Tung, and Kwauk, 1985.)

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559

Figure 45. Instrument for automatic surface tracking and data processing. (Yang, Tung, and Kwauk, 1985.)

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Figure 46. Typical bed collapse curve traced by optical-fiber probe tracking instrument (solids:alumina A66, 140–280 microns). (Yang, Tung, and Kwauk, 1985.)

Not all powders exhibit all the three stages described above, as shown in Fig. 47. Depending on the presence or absence of the above stages, a three-digit qualitative designation has been formulated: 100 Bubble-escape stage only 123 Three stages: bubble escape, hindered sedimentation and solids consolidation 020 Hindered sedimentation only 023 Hindered sedimentation plus solids consolidation

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(1) (,)

(/) ..J

~ Izj E-o cn > cn

t: co

>-

~ co "3 u ~ 0.

(zj ,..j ~ E-o ~ < ~

..,

z o

-fn E-o < ~ 8 ~ ,..j (I.


561

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Figure 47. Essential types of bed collapsing curves.

562 Fluidization, Solids Handling, and Processing Mathematical modeling of the three-stage bed collapsing process led to a more quantitative characterization of powders, in the form of a dimensionless number called the dimensionless subsidence time:

Θ=

d p g  Ze −Zc  ut u2  Z ∞

   

To test the viability of Θ in quantifying fluidizing characteristics, it is plotted against the ratio of incipient bubbling velocity to incipient fluidization velocity, u mb/u mf, the latter being calculated after Geldart (1972). Figure 48 shows that a linear relation exists between ln(u mb/u mf) and θ ¼ as represented by the following empirical relation ln(u mb /u mf) = 4 Θ¼ The straight line starts from Θ = 0 and u mb /u mf = 1, and extends without limit towards the upper right-hand corner. The value of u mb /u mf = 1 obviously signifies bubbling. The corresponding value of Θ is zero, for which stages 1 and 2 of the bed collapsing process take place almost instantaneously. As the fluidizing characteristics improve, the value of u mb/u mf becomes progressively greater than unity, signifying particulate expansion. The corresponding value of Θ also increases, showing a slower bed collapsing process. As the curve tends toward even larger values of Θ, the characteristics of particulate fluidization becomes more predominant. 8.2

Improving Fluidization by Particle Size Adjustment

Catalyst powders with carefully specified particle size distribution have been known to possess good fluidization characteristics. Generally, addition of fine particles to coarse particles tends to improve the latter’s fluidization characteristics. Experiments were thus conducted on binary particle mixtures, each consisting of a fairly close particle size distribution. Figure 49 shows a set of bed collapsing curves for a Geldart Group AA (for Geldart’s classification of solid particles, see Geldart, 1972, 1973) binary solids mixture, two closely sized alumina powders, of average particle diameter 104 and 66 microns, respectively. The curve on the extreme left with 0% fines represents the pure coarse component, which is

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563

barely Group A in fluidizing characteristics, as can be seen from its very brief stage 2. The curve on the extreme right, representing the 100% fine component, demonstrates pronounced Group A, or “123,” fluidizing characteristics with a long stage 2. Curves with intermediate compositions are shown in their ordered locations between the purely coarse and purely fine components.

Figure 48. Plot: u mb /u mf versus Θ¼. (Yang, Tung, and Kwauk, 1985.)

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520

500

e e ~ .. ~ 't: =' "' ] .0 "' -0 := 0 "'

440

~ .t

420

4001

~

~

IiI

0

5

10

II 15

20

25

30

I,s

Figure 49. Bed collapse curves for a group A-A Solid Pair: Solids A104/A66 (Yang, Tung, and Kwauk, 1985).

Figure 50 illustrates the fluidizing characteristics of a set of Group BA mixtures of relatively coarse (212 microns) polyethylene spheres mixed with a relatively fine (27 microns) FCC catalyst. As the fine FCC catalyst is added, there is a progressive predominance of the Group A fluidizing characteristics of the finer component.

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565

480

420

400

Figure 50. Bed collapse curves for a Group B-A solid pair: solids PE201/F27. (Yang, Tung, and Kwauk, 1985.)

Figure 5 1 shows the effect of adding talc (7 microns), Group C, to a coarse sea sand (165 microns), Group B. The beneficial effect of the fines on the fluidizing characteristics is shown to increase to a fines composition of around 25%, and then it diminishes with further fines addition. Figure 52 plots the dimensionless subsidence time 0 for the six sets of Group A-A and Group B-A binary mixtures for different fractions of fines, xr, showing that the improvement of fluidizing characteristics by addition of fine particles increases monotonically with increasing percentage of the tines.


Figure 51. Bed collapse curves for a Group A-C solid pair: solids SA165/T7. (Yang, Tung, and Kwauk, 1985.)

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567

Figure 52. Improving fluidizing characteristics by fines addition, solid groups: A-A and B-A. (Yang, Tung, and Kwauk, 1985.)

Figure 53 is a similar plot, but for the ten sets of Group B-C mixtures. Most curves show that the value of Θ for Group B-C binaries give maxima at certain intermediate compositions. The presence of these maxima suggests that not only fine particles belonging to Group C may improve the fluidizing characteristics of such coarse solids as Group B, but the coarse particles may also improve the fluidizing characteristics of the fine particles, which are known to possess the notorious tendency towards channelling before fluidization sets in. This synergism of Group B-C mixtures testifies to the significance of particle size selection and particle size distribution design, in order to tailor a solid particulate material to certain desired fluidizing characteristics. A mathematical model has been proposed to account for the mutual synergistic action of either particle component on the other in increasing the value of the dimensionless time Θ as shown in Fig. 53, in terms of the mass fraction x2 of fines, and two empirical parameters n 1 and n 2: Θ = Θ1 (1 - x2) (1 + x2n1 ) + Θ2 x2 [1 + (1 - x2) n2 ]

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Figure 53. Improving fluidizing characteristics by fines addition, solid group: B-C. (Yang, Tung, and Kwauk, 1985.)

Mutual synergism of binary mixtures containing fine particles can be quantified in terms of the departure of the Θ-x2 curve from a linear tie line, which signifies absence of synergism, joining Θ1 and Θ2 for the respective components, as shown in Fig. 54. A convenient measure of this departure is the shaded area lying between the Θ-x2-curve and the linear tie line. This can be derived analytically in terms of what will be called the Synergism number Sy, normalized with respect to Θ1+Θ2 of both the coarse and the fine particles:

Sy =

=

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 1  1 1 Θdx − (Θ1 + Θ 2 )  2 Θ1 + Θ 2  0 

∫

1

+

Θ 2 / Θ1

(n1 + 1)(n1 + 2 ) (n 2 + 1)(n 2 + 2 )


569

The larger is the value of Sy, the stronger is the mutual synergistic interaction between the coarse and fine particles. Also, an effective fine particle additive to improve the fluidizing characteristics of coarse particles calls for a large value of Θ m produced with minimal amount of the fine material, that is, a small value of x2. Such modeling has been computerized (Qian and Kwauk, 1986) for on-line data acquisition and analysis for graphing Θ -xf curves directly from the automatic bed-collapsing instrument shown in Fig. 45.

X2

Figure 54. Measure of synergism for binary particle mixtures. (Kwauk, 1984, 1986.)

9.0

WHY BUBBLING AND NOT PARTICULATE FLUIDIZATION

A significant objective of bubbleless gas/solid contacting is to make the aggregative G/S system approach the particulate L/S system for which solid-fluid contacting is far better. Ever since the adoption of fluidization, however, there has been rather inadequate explanation of the vast difference between the bubbling G/S system and the particulate L/S system. The proposition of the so-called “energy-minimized multiscale model” generalizes all particle-fluid systems (Li, Tung and Kwauk, 1988; Li and Kwauk, 1994).

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The Energy-Minimized Multiscale (EMMS) Model

With respect to the highly disparate nature of G/S and L/S fluidization, even at an early stage, the aggregative tendency of L/S fluidization was already discovered when large and/or heavy solid particles were used, and, moreover, clustering of these particles was also reported (Wilhelm and Kwauk, 1948). The particulate phenomenon present in certain G/S systems was noted, however, much later in connection with certain Geldart A-type solids, e.g., the graded, spheroidal FCC catalyst. Aggregative tendency of the normally particulate L/S fluidization on the one hand, and particulate tendency of the normally aggregative G/S fluidization on the other, led to the concept of interpolating these opposite tendencies toward an intermediate regime of transition. Physically the EMMS model consists of a dense phase of clusters dispersed in a dilute phase composed of more or less discrete suspended particles, as shown in Fig. 55 (Li, Tung and Kwauk, 1988). The total fluid flow Ug to the system is divided into two streams, one through the densephase clusters Uc and one through the surrounding dilute phase Uf. Clusters are suspended by the fluid flowing inside them as well as outside them, while the discrete particles in the dilute phase are being entrained by their surrounding fluid. The dense clusters occupy a volume fraction f of the bed volume, the remaining (1 - f) fraction belonging to the dilute phase. Figure 55 also tabulates the symbols for velocity U, for voidage ε and for drag coefficient CD referring to the dense and the dilute phases. There are thus three scales of interaction between the fluid and the particles: a. Microscale of particle size—drag on particle for both the dense and dilute phases b. Mesoscale of cluster size—interaction between dilutephase fluid and clusters c. Macroscale for the overall two-phase flow system, involving its interaction with its boundaries, such as walls and internals.

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dense= c

571

dilute=f

~ I-I

dense

.. el

"I

sup. ftuid vel.

",

sup. ftuid vel.

".,

sup. part. vel.

""

sup. part. vel.

",

sup. re1a. vel.

£,

voidage

£, CDI

voidage drag on part.

C~

drag on part.

dilute x

"J

u,

Figure 55. Physical model for multiscale modelling of particle-fluid system.

According to the above physical model, the following equations have been derived: (i) Force balance for clusters in unit bed volume (ii) Equal pressure drop. As shown in Fig. 55, the dilutephase fluid flow has to support its discrete particles as well as, in part, the clusters in suspension, to yield a pressure drop equal to that of the parallel dense-phase fluid flow (iii) Particle suspension for dilute phase (iv) Continuity for fluid (v) Continuity for particles (vi) Cluster diameter

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572 Fluidization, Solids Handling, and Processing The above six equations for continuity and force balance do not, however, afford a complete description of a heterogeneous particle-fluid system in which a dense phase and a dilute phase coexist. An additional constraint needs to be identified to account for the stability of the system. This constraint is to be provided through the concept of minimal energy. According to this concept, particles in a vertical flow system tend toward certain dynamic array which results in minimal energy. The total energy associated with a flowing particle-fluid system is thus considered to consist of the sum of two portions, one used in suspending and transporting the particles, and one consumed in energy dissipated in particle collision and circulation. Alternately, in terms of the power per unit weight of solids, the total power N is composed of the suspension and transporting portion NST and the energy dissipation portion NED. The former can be split into that for particles suspension NS and transport NT, and in accordance with the multiscale model, also into portions for the dense cluster phase, the surrounding dilute phase and interaction between the two, that is    N ST = N S + NT  N  N ED  

 N dense   N dilute N  inter

Consequently, modeling of a two-phase flow system is subject to both the constraints of the hydrodynamic equations and the constraint of minimizing N. Such modeling is a nonlinear optimization problem. Numerical solution on a computer of this mathematical system yields the eight parameters: ε f, ε c , f, U f, U c, l, U df, U dc plus the various constituent power terms.

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573

Reconciling L/S and G/S Systems

Figure 56 compares the computed results for the FCC/air system against those for the glass/water system, to illustrate the disparate behaviors of G/S and L/S fluidization. The first three insets on the left-hand side show the change of voidages, ε f, ε c and ε, cluster phase fraction f, and slip velocity Us, for the FCC/air system. In the corresponding insets for the L/S system, however, the three voidages ε f, ε c and ε are identical, and the cluster phase fraction f is zero, indicating the absence of clusters throughout the velocity range of Ug, that is, fluidization is homogeneous. Also, the slip velocity Us between solid particles and the surrounding liquid is always less than the terminal velocity Ut of the particles, which remains the asymptotic value for the increasing slip velocity Us as the entering liquid velocity Ug increases. The lowest two insets of Fig. 56 compares the power for suspending and transporting the solid particles NST for the G/S and L/S systems. For the FCC/air system, NST is always less than the total energy N, until it jumps to the latter value at the point of sudden change, while for glass/water, NST is always the same as N in view of its homogeneous nature. Figure 57 shows the gradual transition of the homogeneous glass/ water fluidization to the highly heterogeneous or aggregative glass/air fluidization, as the particle/fluid density ratio ρ p /ρ f increases from water through ethyl ether, and carbon dioxide under different stages of decreasing pressure from its critical condition, to atmospheric air. The appearance and gradual growth of the two-phase structure is evident in the order of the fluids listed. For instance, the curves in the top left inset show that at ρ p/ρ f = 2.55 for glass/water, f = 0 throughout the range of gas velocity Ug, indicating a homogeneous particle-fluid system. When the ratio ρ p/ρ f has increased to 3.19 for glass/ethylether, however, the two-phase structure appears for gas velocities Ug up to 0.4 m/s. This velocity range broadens through carbon dioxide under decreasing pressures from its critical point, until at ρ p/ρ f = 2,162 for glass/air, this two-phase structure has extended beyond Ug = 2 m/s. While Fig. 56 demonstrates, from modeling, the disparate nature between G/S and L/S fluidization, Fig. 57 shows continuity in particle-fluid behavior through properly selected intermediate systems, thus reconciling through theory the phenomenological discrimination between aggregative and particulate fluidization.

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574 Fluidization, Solids Handling, and Processing FCC/airG.=50q/(ml.sl

glass/water

G. =~klf/lm.

.s)

~

0.2

0.4 U.

Figure 56. Comparison between G/S and L/S Systems. (Li, 1987.)

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0.6

0.8


0.5

1.0

1.5

575

2.0

Figure 57. Computation results showing gradual transition from particulate fluidization to aggregative fluidization. (Li, 1987.)

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576 Fluidization, Solids Handling, and Processing 10.0 EPILOGUE Bubbling is the usual phenomenon in most accepted processes of fluidization. It has been demonstrated that species of fluidization with better G/S contacting are available by avoidance of bubbling. This chapter has offered certain tools for discriminating the homogeneous L/S fluidization on the one hand from the bubbling G/S fluidization on the other, in the form of the EMMS model, but this is by no means the only possible method. It has also revealed the significance of particle size choice or design in order to improve the fluidizing characteristics of a given species of fluidization through using the bed collapsing apparatus, but this, too, is by no means the only possible technique. The message of significance is how to “particulatize” a G/S system which is otherwise aggregative.

NOTATIONS A Ax Ar At A´ Bi Cp Cs dp D E f g G h H j kg ks L

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Area; cross-sectional area, cm2 Frontal area of particle, cm2 = d 3p ρ f g∆p/µ 2, Archimedes number, dimensionless Cross-sectional area based on particle terminal velocity, cm2 = A/At, reduced area, dimensionless = hd p /k s, Biot number, dimensionless Heat capacity of gas, cal/mol· oC Heat capacity of solids, cal/gm· oC Particle diameter, cm Diffusivity, cm2/sec = Nu Pr-1/3/Ar or Sh Sc-1/3/Ar, effectiveness factor, dimensionless Friction factor, dimensionless = 980 cm/sec2, acceleration of gravity Mass flow rate of gas, mol/cm2·sec Heat transfer coefficient between particle and fluid, cal/sec·cm2·oC Nonuniformity index, dimensionless = Nu Pr-1/3/Re or Sh Sc-1/3/Re, j-factor, dimensionless Thermal conductivity of gas, cal/cm·sec·oC Thermal conductivity of solid, cal/cm·sec·oC Mass velocity of liquid, gm/cm2·sec


577

m Fixed bed pressure drop exponent, dimensionless n Fluidized bed expansion exponent, dimensionless N = (L/ρ f A)(S/ρ s A), liquid-solids velocity ratio, dimensionless NH Number of heat-transfer stages, dimensionless Nu = hd p /k g, Nusselt number, dimensionless ∆P Pressure drop, gm-wt/cm2 Pr = Cpµ/k g, Prandtl number, dimensionless Re = d p uρ f /µ, Reynolds number, dimensionless S Mass velocity of solids, gm/cm2·sec Sc = µ/ρ f Df, Schmidt number, dimensionless Sh = h D d p /Df , Sherwood number, dimensionless Sy Synergism number, dimensionless T Temperature, oC u Velocity, cm/sec uo Superficial fluid velocity, cm/sec ud Superficial solids velocity, cm/sec u mb Superficial fluid velocity at incipient bubbling, cm/sec u f,umf Superficial fluid velocity at incipient fluidization, cm/sec up Actual velocity of particle, cm/sec us Relative velocity between particle and fluid, cm/sec ut Terminal particle velocity, cm/sec u´ = u /u t , reduced velocity, dimensionless V´ = A´Z´, reduced volume, dimensionless w Particle population density, gm/cm2 xf Weight fraction of fines in binary particle mixture, dimensionless z Distance, cm zi Location of point of inflection for fast fluidization, cm Z Height of fluidized bed, cm Zo Dimensionless distance in accelerative motion of particle; characteristic length for fast fluidization voidage profile, cm Zt Height based on particle terminal velocity, cm Z1 Dimensionless distance in heat transfer z´ = z/Z, reduced distance, dimensionless τ ε ε mf

= GCp /SCs, flowing heat capacity ratio, dimensionless Voidage or void fraction, dimensionless Voidage at incipient fluidization, dimensionless

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578 Fluidization, Solids Handling, and Processing εo ν θ Θ µ ρs ρf ∆ρ Φo

Voidage of fixed bed, dimensionless Fractional heat recovery, dimensionless Time, sec Dimensionless time in particle-fluid heat transfer; dimensionless subsidence time in bed collapsing test Viscosity, gm/cm·sec Density of solid, gm/cm3 Density of fluid, gm/cm3 = ρ s - ρ f, effective density Dimensionless pressure drop in accelerative particle motion

REFERENCES Chen, B., and Kwauk, M., Generalized Fluidization of Non-Ideal Systems, Proc. First Intern. Conf. Circulating Fluidized Bed, 127-132, Halifax, Canada (1985) Coulson, J. M., and Richardson, J. F., “Sedimentation,” Vol. 2, Chemical Engineering, McGraw-Hill, 2nd Ed. (1968); 3rd Ed. (1978) Deng, Y., and Kwauk, M., “Levitation of Discrete Particles in Oscillating Liquids,” Chem. Eng. Sci., 45(2):483–490 (1990) Geldart, D., “The Effect of Particle Size and Size Distribution on the Behavior of Gas-Fluidized Beds,” Powder Technol., 6:201–205 (1972) Geldart, D., “Types of Fluidization,” Powder Technol., 7:285–290 (1973) Houghton, G., “The Behavior of Particles in a Sinusoidal Velocity Field,” Proc. Roy. Soc., A272:33–43 (1963) Houghton, G., “Digital Computer Simulation of Fluttering Lift in the Desert Locust,” Nature 201, pp. 568–570; “Effect of Variation in Lift Coefficient, Phase Angle and Waveform on Fluttering Lift in the Desert Locust,” Nature 202, pp. 870–872; “Simulation of Fluttering Lift in a Bird, Locust, Moth, Fly and Bee,” Nature 202, pp. 1183–1185; “Fluttering Flight Mechanisms in Insects and Birds,” Nature 204, pp. 447–449; “Generalized Hovering-Flight Correlation for Insects,” Nature 204, pp. 666–668 (1964) Houghton, G., “Particle Trajectories and Terminal Velocities in Vertically Oscillating Fluids,” Can. J. Chem. Eng., 44:90-95 (1966) Houghton, G., “Velocity Retardation of Particles in Oscillating Fluids,” Chem. Eng. Sci., 23:287-288 (1968) Kramers, H., “Heat Transfer from Spheres to Flowing Media,” Physica, 12:61 (1946)

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Krantz, W. B., Carley, J. F., and Al-taweel, A. M., “Levittion of Solid Spheres in Pulsating Liquids,” Ind. Eng. Chem. Fund., 12:391-396 (1973) Kwauk, M., “Fluidized Calcination of Aluminum Ore for Pre-Desilication,” (unpubl.), Inst. Chem. Metall., 1964-11–6 (1964c) Kwauk, M., and Tai, D. -W., “Transport Processes in Dilute-Phase Fluidization as Applied to Chemical Metallurgy,” (I). Transport Coefficient and System Pressure Drop as Criteria for Selecting Dilute-Phase Operations; (II). Application of Dilute-Phase Technique to Heat Transfer, (in Chinese, with Eng. abs.), Acta Metallurgica Sinica, 7:264–280; 391–408 (1964) Kwauk, M., “Particulate Fluidization in Chemical Metallurgy,” Scientia Sinica, 16:407 (1973) Kwauk, M., “Fluidized Leaching and Washing,” (in Chinese), Science Press, Beijing (1979a) Kwauk, M., :Fluidized Roasting of Oxidic Chinese Iron Ores,” Scientia Sinica, 22:1265 (1979b); repr. Intern. Chem. Eng., 21:95–115 (1981) Kwauk, M., “The Jigged Reducer,” (unpubl.); (i) 1979; (ii) 1981; (iii) 1983; Inst. Chem. Metall. Kwauk, M., and Wang, Y., Fluidized Leaching and Washing, Chem. E. Symp,. Ser. No. 63, paper D4/BB/1-21 (1981) Kwauk, M., Wang, N., Li, Y. Chen, B., and Shen, Z., “Fast Fluidization at ICM,” Proc. First Intern. Conf. Circulating Fluidized Bed, p. 33–62, Halifax, Canada (1985) Kwauk, M., Fluidization—Idealized and Bubbleless, with Applications, Science Press, Beijing, and Ellis Horwood, U. K. (1992) Kwauk, M., “Fast Fluidization,” Advances in Chemical Engineering, (M. Kwauk, ed.), Vol. 20, Academic Press, U.S.A. (1994) Li, J., Tung, Y., and Kwauk, M., “Fast Fluidization at ICM,” (i) Method of Energy Minimization in Multi-Scale Modeling of Two-Phase Flow; (ii) Energy Transport and Regime Transition in Particle-Fluid Two-Phase Flow; (iii) Axial Voidage Profiles of Fast Fluidized Beds in Different Operating Regions, Second International Conference on Circulating Fluidized Beds, p. 75, 89, 193, Compiegne, France (1988) Li, J., and Kwauk, M., “Particle-Fluid Two-Phase Flow, The Energy-Minimization Multi-Scale Method,” Metallurgical Industry Press, Beijing (1994) Li, Y., and Kwauk, M., “The Dynamics of Fast Fluidization,” Third Intern. Conf. Fluidization, p. 537–544, Henniker, U.S.A. (1980) Liu, D., Liu, J., Li, T., and Kwauk, M., “Shallow-Fluid-Bed Tubular Heat Exchanger,” Fifth Int. Fluidization Conf., p. 401–408, Elsinore, Denmark (1986)

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580 Fluidization, Solids Handling, and Processing Liu, K., “Particle Motion in Standing Waves in Gaseous Medium,” (in Chinese, unpubl.), Inst. Chem. Metall. (1981) Qian, Z., and Kwauk, M., “Computer Application in Characterizing Fluidization by the Bed Collapsing Method,” Tenth Intern. CODATA Conf., Ottawa (1986) Qin, S., and Liu, G., “Application of Optical Fibers to Measurement and Display of Fluidized Systems,” Proc. China-Jpn. Fluidization Symp., p. 258–266, Hangzhou, China (1982) Qin, S., and Liu, G., “Automatic Surface Tracking for Collapsing Fluidized Bed,” Sec. China-Japan Fluidization Symp., Kunming, Elsevier, p. 468 (1985) Reh, L., “Calcination von Aluminumhydroxid in einer zirkulierenden Wirbelschicht,” Chem. Eng.-Tech., 42:447–451 (1970) Reh, L., “Fluidized Bed Processing,” Chem. Eng. Prog. 67:58–63 (1971) Reh, L., “Calcining Aluminum Trihydrate in a Circulating Fluid Bed, a New Technique of High Thermal Efficiency, Metallurges,” Rev. Activ., 1972(15):58–60 (1972) Reh, L., “The Circulating Fluid Bed Reactor—A Key to Efficient Gas-Solid Processing,” Proc. First Intern. Conf. Circulating Fluidized Bed, p. 105– 118, Halifax, Canada (1985) Squires, A. M., “Applications of Fluidized Beds in Coal Technology,” lecture, Intern. School on Heat and Mass Transfer Problems in Future Energy Production, Dubrovnik, Yugoslavia (1975a) Squires, A. M., “Gasification of Coal in High-Velocity Fluidized Beds,” loc. cit. (1975b) Squires, A. M., “The City College Clean Fuels Institute:” Programs for (i) Gasification of Coal in High-Velocity Fluidized Beds; (ii) Hot Gas Cleaning, Symp. Clean Fuels from Coal, I.G.T., Chicago (1975c) Squires, A. M., The Story of Fluid Catalytic Cracking: “The First Circulating Fluid Bed,” Proc. First Intern. Conf. Circulating Fluidized Bed, p. 1–19 (1985) Treybal, R. E., Mass Transfer Operations, McGraw-Hill, 1st Ed. (1955); 2nd Ed. (1968); 3rd Ed. (1980) Tung, Y., and Kwauk, M., Dynamics of Collapsing Fluidized Beds, China-Jpn. Fluidization Symp., Hangzhou, China (1982), Science Press, Beijing, and Gordon Breach, New York, p. 155–166 Tung, Y., Li, J., Zhang, J., and Kwauk, M., “Preliminary Experiments on Radial Voidage Distribution in Fast Fluidization,” Fourth Nat. Conf. Fluidization, Lanzhou, China (1987)

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Tung, Y., Li, J., and Kwauk, M., “Radial Voidage Profiles in a Fast Fluidized Bed,” Proc. of Third China-Jpn. Symp. on Fluidization, p. 139, Beijing, China (1988) Tung, Y., Yang, Z., Xia, Y., Zheng, W., Yang, Y., and Kwauk, M., “Assessing Fluidizing Characteristics of Powders,” Sixth Int. Fluidization Conf., Banff, Canada (1989) Tunstall, E. B., and Houghton, G., “Retardation of Falling Spheres by Hydrodynamic Oscillation,” Chem. Eng. Sci,. 23:1067–1081 (1968) Van Oeveren, R. M., and Houghton, G., “Levitation and Countergravity Motion of Spheres by Nonuniform Hydrodynamic Oscillation,” Chem. Eng. Sci., 26:1958–1961 (1971) Wang, N., Li, Y., Zheng, X., and Kwauk, M., “Voidage Profiling for Fast Fluidization,” First Intern. Conf. Circulating Fluidized Beds, Halifax, Canada (1985) Wilhelm, R. H., and Kwauk, M., “Fluidization of Solid Particles,” Chem. Eng. Prog., 44:201–218 (1948) Yan, Z., Yao, J., and Liu, S., “Turbulence in Grid Zone of Fluidized Reactor,” Proc. China-Jpn. Fluidization Symp., Hangzhou, China, p. 100–111 (1982) Yan, Z., Yao, J. Z., Wang, W. L., Liu, S. J., and Kwauk, M., “Cocurrent Shallow Multistage Fluid-Bed Reactor,” Third Intern. Conf. Fluidization , Kashikojima, Japan, p. 607–614 (1983) Yang, Z., Tung, Y., and Kwauk, M., “Characterizing Fluidization by the Bed Collapsing Method,” Chem. Eng. Commun., 39:217–232 (1985)

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9 Industrial Applications of Three-Phase Fluidization Systems Jack Reese, Ellen M. Silva, Shang-Tian Yang, and Liang-Shih Fan

1.0

INTRODUCTION

Three-phase fluidization systems offer the flexibility of a wide range of operation conditions and contacting modes and are finding increased industrial applications in reacting and physical processing schemes. This is primarily due to the increased use of three-phase fluidization systems in the rapidly growing biotechnology field, but application in other areas abound as well. Successful application of three-phase fluidization systems lies in the comprehensive understanding of the complex transport phenomena mechanisms and associated reaction parameters. For the past 30 years, large research efforts have been put forth to characterize and predict the complex behavior of three-phase systems; however, a complete understanding and the ability to reliably scale up three-phase fluidization systems have yet to be achieved.

582

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Three-Phase Fluidization Systems

583

The term three-phase fluidization, in this chapter, is taken as a system consisting of a gas, liquid, and solid phase, wherein the solid phase is in a non-stationary state, and includes three-phase slurry bubble columns, three-phase fluidized beds, and three-phase flotation columns, but excludes three-phase fixed bed systems. The individual phases in threephase fluidization systems can be reactants, products, catalysts, or inert. For example, in the hydrotreating of light gas oils, the solid phase is catalyst, and the liquid and gas phases are either reactants or products; in the bleaching of paper pulp, the solid phase is both reactant and product, and the gas phase is a reactant while the liquid phase is inert; in anaerobic fermentation, the gas phase results from the biological activity, the liquid phase is product, and the solid is either a biological carrier or the microorganism itself. The important inherent differences in the operating characteristics between fixed bed systems and fluidization systems as given by Fan (1989) are summarized here. Fixed bed systems produce high reactant conversions for reaction kinetics favoring plug flow patterns because of the small axial dispersion of phases and low macromixing present in such reactors. Fixed bed systems offer the advantage of high controllability over product selectivity for complex reactions and low solids attrition and consumption, permitting the use of precious metals as catalysts. In fluidization systems, high macromixing with large axial dispersion is prevalent which produces high reactant conversions for reaction kinetics favoring completely mixed flow patterns. The high mixing in fluidization systems also yields uniformity in the temperature throughout the system and ease in temperature control. The non-stationary nature of the particles in a fluidization system allow for ease in continuous catalyst replacement and, hence, minimum flow maldistribution. Fan (1989) provided a detailed historical development of threephase fluidization systems in reactor applications. Only a brief review of the significant accomplishments and the economic factors affecting the development of three-phase reactors will be provided here. Table 1 provides the important contributions in the application of three-phase fluidization systems for the past several decades. The direct liquefaction of coal to produce liquid fuels was the first commercial reactor application of three-phase fluidization systems, with development having occurred from the mid-1920’s throughout the 1940’s. A large effort was put forth at this time in Europe for the production of liquid fuels from coal as a direct

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584 Fluidization, Solids Handling, and Processing result of the need for such fuels during World War I and II. The commercial production of liquid fuels from direct coal liquefaction peaked at an average annual production of 4.2 million tons of primarily aviation fuel (Donath, 1963). At the end of the war, the direct production of liquid fuels from coal was phased out. Another important process which was developed during and prior to the 1940’s is the application of a three-phase reactor for the reactions known as Fischer-Tropsch (F-T) synthesis. These reactions produce liquid fuels via indirect coal liquefaction involving hydrogen and synthesis gas derived from coal gasification, in the presence of a catalyst. The peak production for the liquid phase F-T synthesis occurred in the early 1950’s at a daily production rate of 11.5 tons of liquid fuels (Kolbel and Ralek, 1980). Fischer-Tropsch synthesis ceased in the mid-1960’s because of the availability of relatively cheap crude oil. However, renewed interest in F-T synthesis in the early 1970’s, caused by the oil shortages of the time, lead to further studies in the commercial scale production of liquid fuels via F-T synthesis in slurry bubble column reactors. This development has continued throughout the 1980’s and into the 1990’s. Table 1. Industrial Development and Application of Three-Phase Fluidization Systems

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Decade

Application

1940's

Coal liquefaction, Fischer-Tropsch synthesis

1950's

Catalytic synthesis of chemicals

1960's

Hydrotreatment of petroleum resids

1970's

Environmental applications

1980's

Biotechnology application

1990's

Biotechnology / Chemical synthesis


585

The development of three-phase industrial-scale reactors in the 1950’s primarily occurred in the chemical process industry through the use of slurry bubble column reactors. In the 1950’s, slurry bubble column reactors were successfully utilized in the chemical process industry for the catalytic synthesis of organic chemicals and polyolefins. The slurry bubble column reactors used for catalytic reactor applications emerged in various designs and were operated under a wide range of flow conditions. Examples of reactor applications in commercial production (Germain et al., 1979) included hydrogenation of glucose to sorbitol, benzene to cyclohexane, butynediol to butenediol, esters to fatty alcohols, aluminum and ethylene to Ziegler alcohol (ALFOL process), and ethylene polymerization (Solvay process). In the 1960’s, following research and development beginning in the late 1950’s, the three-phase fluidized bed reactor was first used commercially for hydrotreating petroleum resids. The technologies developed at this time are known as the H-Oil process and the LC-Fining process and are still currently in commercial operation. The three-phase fluidized bed reactors used in the hydrotreating/hydrocracking operations are commonly referred to as an ebullated bed. The term ebullated bed was first defined by P. W. Garbo in the patent of Johanson (1961) to describe a gasliquid contacting process in contrast to the common industrial term fluidized bed where particles are in fluidization induced by the gas phase alone. Recent developments in the use of ebullated beds for hydrotreating/ hydrocracking include the Texaco T-STAR process, introduced in the early 1990’s (see Sec. 4.2). The development of three-phase reactor technologies in the 1970’s saw renewed interest in the synthetic fuel area due to the energy crisis of 1973. Several processes were developed for direct coal liquefaction using both slurry bubble column reactors (Exxon Donor Solvent process and Solvent Refined Coal process) and three-phase fluidized bed reactors (HCoal process). These processes were again shelved in the early 1980’s due to the low price of petroleum crudes. The 1970’s also brought about increased use of three-phase systems in environmental applications. A three-phase fluidized bed system, known as the Turbulent Bed Contactor, was commercially used in the 1970’s to remove sulfur dioxide and particulates from flue gas generated by coal combustion processes. This wet scrubbing process experienced several

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586 Fluidization, Solids Handling, and Processing operational difficulties including excessive pressure drop, and its use was discontinued in the early 1980’s. Today, however, wet scrubbing systems are still in use for flue gas desulfurization, although the operational mode of these systems is mainly as slurry spray columns rather than as fluidized beds. A second important environmental application of three-phase systems to come out of the 1970’s is the area of biological wastewater treatment. This area of three-phase research has grown rapidly since the 1970’s and continues to grow in the 1990’s with increasing environmental pressures placed on the chemical process industry. Accounts of the historical development of fluidized bed bioreactor application in wastewater treatment can be found in Fan (1989), and Heijnen et al. (1989). The 1980’s and the early 1990’s have seen the blossoming development of the biotechnology field. Three-phase fluidized bed bioreactors have become an essential element in the commercialization of processes to yield products and treat wastewater via biological mechanisms. Fluidized bed bioreactors have been applied in the areas of wastewater treatment, discussed previously, fermentation, and cell culture. The large scale application of three-phase fluidized bed or slurry bubble column fermentors are represented by ethanol production in a 10,000 liter fermentor (Samejima et al., 1984), penicillin production in a 200 liter fermentor (Endo et al., 1986), and the production of monoclonal antibodies in a 1,000 liter slurry bubble column bioreactor (Birch et al., 1985). Fan (1989) provides a complete review of biological applications of threephase fluidized beds up to 1989. Part II of this chapter covers the recent developments in three-phase fluidized bed bioreactor technology. The design and scale-up of three-phase fluidization systems is a challenging endeavor requiring a sound knowledge of all facets of reactor engineering and the underlying chemical engineering technologies (Tarmy and Coulaloglou, 1992). Research efforts in the 1990’s, using state-ofthe-art imaging techniques, have begun to look inside small scale threephase systems to better understand the fluid mechanics and phase interactions occurring in the system which can lead to better predictive tools for the design of three-phase reactor systems. While these studies provide insight into the hydrodynamics of the system, an important area in the industrial application of three-phase reactor systems which requires further study is the coupling of the hydrodynamics and the reaction kinetics. The complex reaction mechanisms occurring at all scales in three-phase

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reacting systems need to be studied in relation to the hydrodynamic behavior transpiring at the various scales. These areas of study require improvements in experimental techniques and instrumentation to provide the necessary data required in the verification of any predictive models. New applications and novel reactor configurations or operational modes for three-phase systems are continually being reported. These include the operation of a three-phase fluidized bed in a circulatory mode (Liang et al., 1995), similar to the commonly applied gas-solid circulating fluidized bed; the application of a three-phase fluidized bed electrode that can be used as a fuel cell (Tanaka et al., 1990); magnetically stabilized three-phase fluidized beds; centrifugal three-phase reactors; and airlift reactors. This chapter provides an overview of recent and nontraditional industrial applications of three-phase fluidization systems. Fan (1989) has provided comprehensive tables on examples of three-phase fluidization applications as well as several chapters on common industrial applications, and readers are referred there for further information. Shah (1979) and Deckwer (1992) also provided industrial examples of three-phase fluidization systems as well as details on reactor engineering and transport phenomena of these systems. The current chapter is divided into two parts, the first focusing on chemical process applications, and the second considering the biological applications of three-phase fluidization systems. In Part I, two nontraditional industrial applications of three-phase systems are highlighted; these are smelting operations and operations in the processing of pulp and paper, including cooking of wood chips to produce pulp and flotation for recycling paper. An overview of recent developments in several classic applications of three-phase fluidization systems in the chemical process industries, including hydrotreating/hydrocracking using ebullated bed technology and hydrocarbon synthesis using slurry bubble column, are also presented in Part I. Part II of this chapter delineates the recent progress in the application of three-phase fluidization systems to the rapidly expanding biotechnology area. The current status of three-phase bioreactor research, the important differences between three-phase chemical and biological fluidized systems, and design considerations for three-phase bioreactors are covered.

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Part I: Smelting Reduction, Paper Processing, and Chemical Processing 2.0

SMELTING REDUCTION

2.1

Introduction

Smelting reduction of iron ore is an intensive area of research for the replacement of the classical ironmaking process (coke ovens, agglomeration, blast furnaces). Smelting reduction has been defined in several ways; however, the most common definition of these processes in the recent literature is the production of hot metal from iron ore without the use of metallurgical coke (Fine et al., 1989; Oeters et al., 1994). Smelting reduction offers several advantages over blast furnace operation, which are listed in Table 2, but the process may still require another five to ten years before making a significant impact on the worldwide production of iron, primarily because the economics of the smelting reduction process are still uncertain. The key characteristics of smelting reduction as an alternative to the older blast furnace process is the ability to directly use iron ore and coal, which eliminates the need for preprocessing the iron ore and is significant for countries such as the United States that have large reservoirs of coal and the ability to handle an increased load of scrap metal, which is consistent with the worldwide environmental trend of reuse, reduce, and recycle. The development of various smelting reduction processes is occurring simultaneously worldwide. Table 3 summarizes the smelting reduction processes under development and in production around the world. Smelting reduction processes can be classified into melter-gasifiers and iron-bath reactors (Oeters et al., 1994). In the melter-gasifier process, the coal is combusted to carbon monoxide (CO) and hydrogen (H2) to provide the heat to melt iron pellets previously reduced from iron oxide pellets. The iron oxide pellets are reduced via direct reduction using the off-gas from the combusted coal. The most fully developed smelting reduction process is a melter-gasifier known as the Corex process, which is in operation in South Africa (Flickenschild, 1991). The iron-bath

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reactors, however, seem to be the smelting reduction process of choice for current and future applications. Iron-bath reactors consist of three layers or zones—at the bottom of the reactor is the liquid metal bath, which contains the reduced iron product; on top of the liquid metal is the slag layer, which contains all of the solid material in the reactor, including the iron ore and coal, and is where the combustion of the coal and reduction of the iron ore occurs; the upper portion of the reactor is primarily a gaseous region where secondary or post-combustion of the gases, produced during the initial combustion of the coal in the slag layer, occurs. In iron-bath reactors, final reduction of the iron ore and post-combustion of coal occur simultaneously, while the off-gas from the post-combustion is used for the pre-reduction of iron ore in a separate reactor, generally a gas-solid fluidized bed or circulating fluidized bed. The focus of this section will be on iron-bath reactors, which are unique three-phase systems consisting of molten metal as the liquid phase and iron ore, coal, and slag particles as the solid phase. The gas phase consists of bottom injected oxygen, air or nitrogen, used for agitation of the molten metal bath, and top injected oxygen for post-combustion. Table 2. Advantages of Smelting Reduction over Coke Oven and Blast Furnace · Elimination of coke oven by direct use of coal · Direct use of fine ores, do not require agglomeration plant · Lower investment cost as a result of above factors · Lower operating costs due to cheaper raw materials and lack of preparation plants · Lower level of emissions · Higher specific productivity · Increased flexibility in production and operation · Improved controllability · Possibility for direct alloying of steel melts

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590 Fluidization, Solids Handling, and Processing Table 3. Smelting Reduction Processes Under Development Throughout the World Country

Smelting reduction process

Japan

DIOS (Direct iron ore smelting)

Australia

HIsmelt (High intensity smelting)

USA

AISI (Direct steelmaking program)

Germany

Corex

France and Germany

Jupiter

Great Britain, Holland, Italy

CCF (Cyclone converter furnace)

Canada

LB (Lu-Bryk process)

Russia

ROMELT (Rossia Melt)

2.2

Principles of Smelting Reduction

Smelting reduction in iron-bath reactors consists of two major steps, pre-reduction of the iron ore in a fluidized bed reactor or prereduction reactor (PRR) and smelting reduction in a smelting reduction furnace (SRF), see Fig. 1. A brief summary of the components and chemical reactions in the SRF are given in Fig. 2. The coarse coal particles are charged directly into the furnace, while lump and fine iron ore are injected into the iron-carbon slag layer or simply added to the top of the vessel at an approximate temperature of 800°C after pre-reduction by the off-gas. The iron ore is reduced at temperatures between 1400–1600°C. The gaseous products from the combustion of coal and reduction of iron oxide contain large amounts of carbon monoxide and hydrogen. A layer of slag on the liquid iron bath protects the liquid iron product from reoxidation. The predominant components of the slag layer are lime, silica, and iron oxide. On the surface of the slag and in the upper portion of the furnace, the burning of the carbon monoxide and hydrogen with oxygen and/or air injected through a top penetrating lance is called post-combustion. The

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heat generated by post-combustion, if efficiently transferred back to the iron bath, significantly reduces coal consumption. In the SRF, the iron bath is also agitated by bottom injection of nitrogen, oxygen, carbon dioxide, or air to improve heat transfer and increase slag circulation. The hot metal is continuously removed from the bottom of the SRF at temperatures around 1500°C.

~

Processgas

Prereduction Coal ~:

~ ~ ~.

Melting Reduction

,, '. ., : .:

Slag

Iron melt ~=~

Figure 1. Two-stage smelting reduction process using an iron-bath reactor. (From Japan Metal Bull., 1988.)

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Coal

Exhaust gas

Prereduced ore

,..t--+

(

FeO Fe203

Fe

Gas

,

,.-

I

,,

/,1

~J::::::==-Post

I--r--

.H2

Combustion CO

)

(

+ O2..

C02

1 \ \ ,-/

It\ 10 2 \ ,-/

CxHy

H2Q

I

Carbon

gasification

C + tO2..CO

t

Smelting reduction FeO + C-Fe + CO t

Heat transfer

Coal decomposition Ore melting

t

Stirringgas

Figure 2. Components and chemical reactions of an iron-bath smelting reduction furnace. (From Shinotake and Takamoto, 1993.)

In a two-stage smelting reduction process, several primary factors influence the operation, including the degree of pre-reduction, degree of post-combustion, slag behavior, heat transfer efficiency, and bath agitation. During the pre-reduction of the iron ore using the off-gas, the carbon monoxide requirement for pre-reduction and the evolution of carbon monoxide in the smelting reduction process can be harmonized (Oeters et al., 1994). The two-stage smelting reduction process can be optimized at a certain degree of pre-reduction depending on the CO requirement for pre-reduction and the evolution of CO in smelting reduction. A reasonably high degree of post combustion associated with a high heat transfer efficiency is also essential for the efficient operation of a smelting reduction process. The complex behavior of the multiphase slag layer plays an integral role in the performance of the furnace because of the layer’s importance in heat transfer, mass transfer, and chemical reaction. The following sections review current research and factors affecting postcombustion and slag behavior.

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Three-Phase Fluidization Systems 2.3

593

Post-Combustion and Heat Transfer in SRF

The energy required in smelting reduction in iron-bath reactors for heating the bath, for melting scrap material, for reducing iron oxide, and for gasifying coal is supplied primarily by the combustion of carbon or coal to carbon monoxide (Zhang and Oeters, 1991a). The total energy available in the coal is not entirely utilized because of incomplete combustion of the coal. Therefore, to improve the energy efficiency in the furnace, carbon monoxide and hydrogen produced from the combustion of coal in the slag layer undergo further combustion to carbon dioxide and steam in the upper portion of the furnace providing additional heat that can be transferred back into the slag layer and iron-bath. A measure of the post-combustion is the post-combustion ratio (PCR), which is one of the important process parameters in smelting reduction processes, defined by:

Eq. (1)

PCR =

% CO2 + % H2 O %CO + % CO2 + % H2 + % H2 O

Higher post-combustion ratios imply that more heat is generated which can be utilized in carrying out the smelting operations. Typical postcombustion ratios are in the range of 40–60% (Brotzmann, 1989; Ibaraki et al., 1990; Romenets, 1990; Takahashi et al., 1992; Katayama et al., 1993b). As in any process, however, not all the heat generated by the postcombustion reactions is transferred back to the bath; thus a second parameter known as the heat transfer efficiency (HTE) is defined as the portion of the heat generated by post-combustion transferred to the bath.

Eq. (2)

  Excessheat in the off -gas  ×100 HTE =  1 Heat generated by post combustion  

With this definition, a heat transfer efficiency of 100% implies that the temperature of the off-gas will be the same as the temperature of the bath. The HTE reported in the literature are in the 80–90% range (Ibaraki et al., 1990; Takahashi et al., 1992; Katayama et al., 1993b). Several authors (i.e., Gou et al., 1993, and Gudenau et al., 1993) have indicated that this definition has limitations because the heat losses to the furnace walls

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594 Fluidization, Solids Handling, and Processing (refractories) and the top penetrating gas lance are not taken into account. The losses are accounted for in the HTE definition proposed by Gudenau et al. (1993) given in Eq. (3).

Eq. (3)

 heat loss of the off - gas + other losses in furance  HTE = 1   total heat generated by post combustion  

Several studies have investigated the relationship between PCR and HTE with the results, in general, demonstrating an inverse relationship (Zhang and Oeters, 1991b; Farrand et al., 1992; Hirata et al., 1992; Katayama et al., 1992; Gudenau et al., 1993). Ideally, high levels of PCR coupled with high levels of HTE are desirable; however, to achieve high levels of PCR, the amount of oxygen fed to the reactor must increase. This can lead to reoxidation of the iron droplets in the slag layer and may increase the off-gas temperatures to the point of damaging the refractory material of the furnace. The increase in the off-gas temperature can be minimized with an increase in the heat transfer to the slag layer; however, the mechanisms of heat transfer may limit the degree of post-combustion. The factors affecting the relationship between PCR and HTE are numerous and complicated (including amount and contents of coal); further studies in theoretical analysis and experimental evaluation are required. The main aspects of heat transfer in a SRF are the heat transfer from the post-combustion gas to the slag layer and from the slag layer to the iron bath. In the mixing process of the slag layer, slag and iron particles are ejected into the gas space where they acquire heat from the post-combustion gas. The particles, after a short time in the gas space, fall back into the slag layer with the iron contained in the particles melting due to a rise in the temperature of the particles, and accumulating in the molten iron bath. The heat is transferred in this way from the post-combustion gases in the upper portion of the reactor to the liquid iron bath at the lower portion of the reactor. Heat transfer from the gas phase to particulate phase consists of convective and radiative transfer. Oeters et al. (1994) stated that, for conditions where the post-combustion took place by the top blown gas penetrating the slag layer, the main resistance to heat transfer was between the gas phase and slag, and that the radiative transfer between the gas and slag was the most important heat transfer mechanism. Katayama et al. (1992) reported that under actual operating conditions, 30% of the heat

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transfer could be explained by radiation and convection, while the rest was attributed to circulation of superheated carbonaceous materials within the slag. Ibaraki (1994) considered the heat transfer to be caused by two mechanisms; (i) radiation heat transfer from the gasses in the upper portion of the reactor to the slag layer and (ii) heat transfer from the flame of the oxygen jet through the surface of the slag cavity created by the oxygen jet. For the first mechanism, the radiation increased the temperature of the slag surface, and the slag motion made the temperature homogenous throughout the slag layer. Ibaraki (1994) concluded, based on his prediction, that the radiation heat transfer is not significant, accounting for 20% of the heat transfer. Based on another calculation using a flame temperature of 2600 K and an estimation of the cavity surface area, Ibaraki (1994) stated that the second mechanism accounted for only 20% of the heat transfer. Since these two mechanisms accounted for only 40% of the total heat transfer, Ibaraki (1994) proposed a third mechanism for heat transfer that accounted for the final 60% of the heat transfer. The third heat transfer mechanism is the transfer of heat from the combustion process in the slag to the reduction sites. The oxygen from the jet is not completely reacted upon reaching the slag layer, therefore, a portion of the oxygen reacted with gas, iron droplets, and char in the slag. In order to provide further insight into the post-combustion ratio and the heat transfer efficiency, the factors that affect the PCR and HTE will be delineated. The factors that affect the PCR and HTE will be discussed separately with the understanding that a complex relationship may exist between the two parameters. The factors that affect the PCR are shown in Table 4, and Fig. 3 demonstrates the primary conditions for postcombustion. The PCR should be kept relatively high, since the fuel consumption decreases with an increase in the PCR at the same HTE (Aukrust, 1993). However, as mentioned, high PCR may lead to problems due to increases in: - Off-gas temperature, with detrimental effects on the refractories and gas handling systems (Fruehan et al., 1989) - Oxidation potential of the off-gas (Brotzmann, 1989) - Volume of the off-gas, which may cause difficulties with the gas and dust handling (Hoffman, 1991)

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596 Fluidization, Solids Handling, and Processing Since a detailed discussion of the factors listed in Table 4 is beyond the scope of this review, the reader is referred to the listed references for further information. Several of the factors affecting the PCR also affect the slag layer and these are discussed in more detail in Sec. 2.4. The heat transfer efficiency is significantly affected by the slag layer properties and behavior; therefore, those factors other than slag phenomena that affect HTE are presented in Table 5. As for the PCR, it is desirable to keep the HTE as high as possible. An increase in HTE at the same PCR decreases fuel consumption (Fruehan et al., 1989; Keogh et al., 1991) and the off-gas temperature (Fruehan et al., 1989; Takahashi et al., 1992).

Table 4. Factors That Affect the PCR

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Factor

References

· Carbonaceous material - volatile matter, coal feed rate, location of coal injection

Keogh et al. (1991); Hardie et al. (1993); Ibaraki (1994)

· Ore feed rate

Shinotake and Takamoto (1993)

· Iron droplet behavior - mixing, reoxidation

Ibaraki et al. (1990); Zhang and Oeters (1991b); Ibaraki et al. (1995a)

· Slag layer foaming

Tsujino et al. (1985); Hirai et al. (1987); Farrand et al. (1992); Ibaraki et al. (1995a)

· Temperature and depth of the metal bath

Keogh et al. (1991); Hardie et al. (1993)

· Agitation of metal bath

Ibaraki et al. (1990); Hirata et al. (1991); Katayama et al. (1992)

· Type of agitation gas and location of injection

Ibaraki et al. (1990); Hirata et al. (1991); Hardie et al. (1993)

· Lance design - height, number of holes

Takashiba et al. (1989); Ibaraki et al. (1990); Farrand et al. (1992); Takahashi et al. (1992);

· Type of top-blown injector gas oxygen, air

Keogh et al. (1991); Zhang and Oeters (1993a,b)

· Furnace pressure

Takahashi et al. (1992)


597

Figure 3. Key factors to achieve efficient post-combustion in an iron-bath smelting reduction furnace. (From Takahashi et al., 1992.)

Table 5. Factors Other Than Slag Phenomena That Affect the HTE Factor

References

· Volatility and feed rate of coal

Ibaraki et al. (1990); Takahashi et al. (1992)

· Ore feed rate


· Reoxidation of metal droplets above the bath

Ibaraki et al. (1990); Hirata et al. (1991); Keogh et al. (1991);

· Metal bath agitation

Tanabe et al. (1989); Ibaraki et al. (1990); Hirata et al. (1991); Katayama et al. (1992)

· Lance height

Tanabe et al. (1989); Gou et al. (1993); Takahashi et al. (1992)

· Type of top blown injector gas oxygen, air

Gudenau et al. (1993)

· Furnace pressure


· Heat losses due to dust generation and furnace geometry

Gou et al. (1993); Ibaraki et al. (1995b)

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598 Fluidization, Solids Handling, and Processing Several authors have attempted theoretical modeling of post-combustion (Fruehan et al., 1989; Zhang and Oeters, 1991a, 1991b, 1993a, 1993b; Hardie et al., 1992; Gou et al., 1993; Gudenau et al., 1993). The models consider the convective and radiative nature of the heat transfer under steady-state conditions for various gas types including oxygen and air. The model equations consist of heat and mass balances of the various components in the furnace (i.e., iron and carbon dioxide) and, in some cases, momentum balances for the flow behavior. Gou et al. (1993) provided one of the more comprehensive models for heat transfer. The model, by incorporating the momentum balances, allowed not only for the calculation of the composition and temperatures but also the steady-state gaseous flow patterns. The model determined the location, shape, and temperature of the combustion flame front and coupled equations for the turbulent convective transfer to those for the radiative transfer. They use the DeMarco-Lockwood flux model to calculate the radiation transfer for a turbulent, high-temperature reacting flow system. In the DeMarcoLockwood flux model, the radiation fluxes for axisymmetric flow are expressed in the following form:

Eq. (4)

∂ 1 ∂Fz 4 = ka( 2 F z - F r - σ T 4 ) ∂z k a ∂z 3

Eq. (5)

1 ∂ r ∂ Fr 4 = ka( 2 F z - F z - σ T 4 ) r ∂z k a ∂r 3

where Fz and Fr are the radiation heat fluxes in the z and r directions, respectively. The radiation heat fluxes are coupled to the convective heat transfer via a source term in the transport equation for the stagnation enthalpy of the form:

Eq. (6)

SH =

16 9

k a ( F z + F r - 2σ T 4 )

where S H is the source term from the enthalpy equation.

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Future requirements for the modeling of smelting furnaces will be for a detailed account of the heat transfer mechanisms as developed by Zhang and Oeters (1991a, 1991b, 1993a, 1993b) and a consideration of the fluid dynamics of the system such as that described above by Gou et al. (1993) while considering the effects of the raw materials to maintain model flexibility. The slag layer has tremendous effects on the overall behavior of the furnace, but the complex multiphase nature of the system is still in the early development stages of modeling. Another important factor that needs to be researched further and included in models is the reduction kinetics of iron oxide from liquid slags (Fine et al., 1989; Oeters, 1989; Oeters and Xie, 1995). 2.4

Slag Layer Behavior

The slag layer is composed of molten slag, solid materials such as carbonaceous materials and undissolved ore, gas bubbles, and metal droplets. The operational characteristics, including post-combustion, heat efficiency, dust generation, and carbon content of the bath, of a SRF are highly influenced by the behavior of the slag layer (Katayama et al., 1993a). The slag added to a SRF consists primarily of lime (CaO) and silica (SiO 2) with significant amounts of iron oxide (FeO) becoming present when the smelting reduction process commences. Small amounts of additives, such as P2O5, S, MgO, or CaF2, may also be present in the slag layer. The slag layer is present primarily to serve as a heat flux between the combustion space and the metal bath and to protect the metal bath from the oxidizing atmosphere above the slag. The two most researched topics concerning the slag layer behavior are the optimal height of the slag layer and the mechanisms and control of foaming of the slag layer. Smelting reduction processes are conducted with both a relatively thin slag layer, as in the HIsmelt process developed in Australia (Hardie et al., 1992), and with a thick slag layer, as in the DIOS process developed in Japan (Inatani, 1991). The main reactions in smelting reduction processes are the gasification of coal and iron oxide reduction in the slag. These reactions produce large quantities of gaseous carbon monoxide and hydrogen on the slag/ iron or slag/carbon surface. The slag foam is created by the gas bubbles formed as a result of these reactions becoming entrained in the slag, and in not being able to readily coalesce, forming a system of tightly packed bubbles separated by one another by thin films of liquid slag (Gaskell,

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600 Fluidization, Solids Handling, and Processing 1989). In studies conducted on the foaming behavior of metallurgical slags, it is generally agreed that the principle factors in controlling the foaming behavior are surface tension, viscosity, viscoelastic characteristics of the liquid films, basicity, solid particles, and bubble size. There is, however, considerable disagreement in the relative importance of each factor (Gaskell, 1989; Utigard and Zamalloa, 1993). Obviously, in a complicated three-phase system, it becomes extremely difficult to ascertain the individual effects of the stated factors, which is why many studies on the foaming behavior are conducted in non-reacting systems. The stabilization and intensification of foams are favored for low surface tension and high viscosity of the slag (Fine et al., 1989). Gaskell (1989) provided an example of how the presence of a surface active solute, in this case SiO 2, in a nominally-basic slag led to stabilization of the foam. The surface active solute decreased the surface tension of the slag by preferential segregation of SiO 2 to the slag surface and increased the viscosity as a result of the surface being more siliceous than the bulk melt (i.e., the viscosity of the surface is higher than that of the bulk melt). Aukrust and Dowling (1991) reported that a foaming slag was stable when the basicity was between 1.2 and 1.5. Good summaries of the more fundamental behavior of slag properties can be found in Gaskell (1989), Gudenau et al. (1992), Hara and Ogino (1992), Ogawa et al. (1993), Utigard and Zamalloa (1993), and Gou et al. (1994). Other factors that are more engineering or design oriented are also found to affect the foaming behavior of the slag layer. Gudenau et al. (1992) reported that the top oxygen blowing parameters, such as lance penetration distance, gas velocity, and expansion angle of the gas stream, affected the foam generation and proper selection of these parameters can assist in controlling excess foaming. They also found that the foaming decreased with increased pre-reduction degree. Ibaraki et al. (1990) found that weak bath agitation and a low bath temperature promoted foaming. They also reported that bath agitation affected the iron droplet mixing in the slag and the slag circulation. An increase in bath agitation lifted more iron droplets into the slag resulting in lower PCR and increased the ferrous dust in the off-gas. They suggested a bath agitation of 2 to 4 kW/ton of metal. Ogawa et al. (1992) reported that an increase in iron ore feed rate promoted foaming.

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Foaming of the slag layer is inevitably going to occur during the smelting reduction process, therefore, ways of controlling the foaming to prevent overflow of the system have been studied. The most widely reported and possibly easiest way to control foaming is the addition of carbonaceous material to the slag. The DIOS smelting reduction furnace, under development in Japan, is chosen to demonstrate the control of the slag layer foaming because the process has been described in the literature in more detail than other processes and it operates with a thick layer of slag. The thick layer of slag is required to shield the stirred metal bath from the top blown oxygen jet. The thick slag layer consists of molten slag, carbonaceous materials, and gas bubbles. The carbonaceous material, consisting of char or coke, is suspended in the slag to control foaming. Ogawa et al. (1992) reported that the carbonaceous material promoted the coalescence of small bubbles into larger bubbles that are able to rise through the slag layer. With the use of x-ray fluoroscopic observation of the slag layer in a 1-ton furnace, Ogawa et al. (1992) concluded that the carbonaceous material controlled the slag foaming by increasing the coalescence of small bubbles on the surface of the carbonaceous material because the carbonaceous material was not easily wettable with slag and, therefore, the CO gas was likely to spread along its surface. The amount of carbonaceous material required for stable operation is between 10–20% of the slag’s weight (Ibaraki et al., 1990; Ogawa et al., 1992). Katayama et al. (1993a) measured the slag height and by assuming that all the carbonaceous materials were entrapped in the slag layer, calculated the volume proportions of the slag layer to be as follows: molten slag: 22–28 vol%, carbonaceous materials: 22–28 vol%, and gas bubbles: 44–55 vol%. Lump carbonaceous material is preferred over fine material, because the turbulence created by the larger particles lead to more frequent contact with the small bubbles, and the fine material is only sparsely scattered when charged into the furnace (Ogawa et al., 1992). The slag layer plays important roles in heat transfer and the reduction of iron oxide. The role of the slag layer in heat transfer was described in Sec. 2.3. Katayama et al. (1992) proposed that the reduction of iron oxide occurred at the following three locations (i) Interface between slag and metal bath (ii) Interface between slag and metal droplets (iii) Interface between slag and carbonaceous materials.

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602 Fluidization, Solids Handling, and Processing They estimated the overall rate of reduction to be equal to the rate of addition of ore at stable operation. By determining the overall rate as a function of the amount of slag in both 5-ton and 100-ton furnaces, they calculated the amount of reduction at each site. Katayama et al. (1992) found that for the 5-ton furnace operation, a high proportion (70%) of the reduction occurred at the bulk metal slag interface. The proportions of reduction at locations 2 and 3 were found to increase with increasing experimental scale and the amount of slag. For the 100-ton furnace containing a large amount of carbonaceous material, the amount of reduction at each site was almost equal (34% for site 1, 33% for sites 2 and 3). Ibaraki et al. (1990) reported that the reduction reaction took place in the slag bulk as well as in the condensed iron zone around the static slag/ metal interface. They also reported that the reduction kinetics decreased as the oxygen jet penetrated the iron bath and concluded that the oxygen jet in the slag not only interfered with the reduction reaction, but also reduced the amount of post-combustion and increased the ferrous dust content in the off-gas. They recommended that the blown oxygen should not reach the condensed iron zone, located in the lower one-third of the slag. Katayama et al. (1993a) reported that the temperature of the lower 75% of the slag layer was almost equal to that of the metal bath, while the upper 25% could have large fluctuations in the temperature caused by the interaction with the gas phase in the upper portion of the furnace where temperatures may be 200–300°C higher than in the metal bath. As mentioned previously, the modeling of the slag behavior is an area that requires further development. Ogawa et al. (1993) modeled the physical behavior of foaming slags by a mechanistic model that considered the effects of the physical properties of the slag and metal on the foam height. The model demonstrated that not only do the surface tension and viscosity of the slag affect the foam height, but the slag/metal interfacial tension and the surface tension of the metal affect the foam height as well. Gou et al. (1994) applied a one-dimensional, fluid mechanic model for gas and liquid two-phase flow to calculate the foam height and void fraction variation with gas velocity. The model requires an experimentally determined parameter of the ratio of the drag coefficient to the effective bubble diameter. The model was able to predict the slag level of those experiments that were used in determination of the above stated parameter. The model does not directly account for the raw materials.

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603

Future of Smelting Reduction of Iron Ore

Table 6 provides areas of smelting reduction that require further research. The economics of the current smelting reduction processes under development need to be more accurately determined as the development stages become larger. One estimation of the economics of smelting reduction and a comparison between blast furnace, direct reduction, and smelting reduction was provided by Steffen (1989). As seen from Table 3, the development of smelting reduction processes is occurring worldwide. Ibaraki et al. (1995a, 1995b) summarized research performed in Japan between 1988 to 1991 using a 100 metric ton smelter, indicating the highest production rate achieved was 36.4 metric tons of iron/hour. Fine et al. (1989) estimated that based on research and development efforts in Japan, the USA, and Europe, the implementation of a smelting reduction process to produce steel is possible by the year 2000. Table 6. Areas of Smelting Reduction Requiring Further Research (Fine et al., 1989; Oeters, 1989) · Melting and reduction mechanisms of pre-reduced ore in iron oxide containing slag and in carbon containing iron · Sticking behavior of ore fines · Kinetics of high temperature oxidation of liquid iron · Mechanism of carbon monoxide formation under the conditions of smelting reduction · Flow processes in thoroughly mixed metal-slag-gas systems · Mass and heat transfer during post-combustion in thoroughly mixed metalslag-gas systems · Reactions of FeO-bearing melts with refractory materials · Optimization of reactor geometry in terms of charging and removal of materials · Optimization of slag height and metal bath agitation · Alternative heat sources such as electrical energy · Use of increased amounts of scrap metal and lower grades of iron ore · Removal of sulfur and other trace elements from the off-gas and liquid iron

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PAPER PROCESSING

3.1

Introduction

The pulp and paper industry is characterized by many operations that are performed under multiphase flow conditions. Typical examples of the major processes that involve multiphase phenomena are the cooking of wood chips in a chemical solution to remove unwanted cross-linking components (lignin) leaving behind a fibrous wood pulp; bleaching of pulp in three-phase mixtures of pulp slurry and bleaching gas (chlorine, chlorine dioxide, oxygen, or ozone); and de-inking of recycled paper in flotation-type devices. Each of these processes and the nature of their three-phase behavior will be discussed in this section. The operational characteristics of these processes may differ somewhat from traditional three-phase fluidized beds, nevertheless, the processes still consist of three-phase mixtures that require intimate contacting of the phases. Since the pulp and paper mills of the future may only consist of improving on existing technologies (Rickard, 1994; Meadows, 1995), a further understanding of the three-phase nature of these processes will assist in attaining these necessary process improvements. Three-phase mixtures in pulp and paper are unique because of the complex behavior of the solid phase. The solid phase (wood chips or paper pulp) in the pulp and paper industry is a low-density fibrous material (primary cellulose) of varying size, shape, and thickness. The nature of the solid phase is complicated because of the ability of the pulp fibers to absorb water and swell to several times the original volume of the oven dried (O. D.) pulp. This absorbed water becomes part of the fibers, and is not considered part of the suspending medium (Stenuf and Unbehend, 1986). This type of solid phase is in contrast to the high density, primary spherical solid phase seen previously in the smelting operation and the catalyst material used in chemical or hydrocarbon processing. While twophase pulp slurry pipe flow has been extensively studied (Stenuf and Unbehend, 1986), highlighted by the demonstration of the non-Newtonian and drag reduction characteristics, the three-phase flow behavior requires further study (Lindsay et al., 1995; Reese et al., 1996).

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605

Chemical Pulping of Wood Chips

Pulping of wood chips refers to the process by which the wood chips are degraded into fibers by removal of the lignin from the cellulose and hemicelluloses. The lignin is a highly polymerized substance that holds the cellulose fibers together. The two primary methods of pulping are chemical and mechanical. Chemical pulping consists of cooking the wood chips with the appropriate chemicals in an aqueous solution at elevated temperature and pressure. The chemicals used in chemical pulping can be either alkaline (kraft process) or acidic (sulfite process). The alkaline kraft process has become the chemical pulping method of choice because of advantages in chemical recovery and pulp strength. However, the formation of organic sulfides in the kraft process has caused environmental concern due to their escape as malodorous gases. A brief overview of the kraft process will be given here. Details on mechanical pulping and the sulfite process can be found in Smook (1992). In the kraft process, an aqueous mixture of sodium hydroxide (NaOH) and sodium sulfide (Na2S) is used to break the lignin away from the cellulose fibers. The chemical reactions that occur are complex and not completely understood. The alkaline attack essentially breaks down the lignin molecules into smaller fragments whose sodium salts are soluble in the solution. The cooking of the wood chips is completed when the kappa number, a measure of the residual lignin content in the pulp, of the pulp has achieved a desirable level. The primary objective in the cooking of wood chips is to consistently obtain the desired kappa number with a minimal amount of cellulose degradation and a minimal level of unusable (reject) material, producing a pulp of uniform quality while working with unavoidable variations in the wood feedstock. The two factors that drive the kraft pulping reactions are alkali concentration and temperature. The rate of removal of the lignin (delignification) more than doubles over the range of normal cooking temperatures (155–175°C) with every 10°C increase in temperature (Smook, 1992). One drawback to the kraft process (and chemical pulping in general) is that while the method is able to remove a large portion of the lignin, a certain amount of degradation of the cellulose will also occur. Typical yields from kraft processes are between 40% and 50% of the original wood substance compared with 85% to 95% from mechanical pulping. Various factors that affect the chemical cooking of wood chips via the kraft process are listed in Table 7.

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606 Fluidization, Solids Handling, and Processing Table 7. Factors Affecting Kraft Cooking (After Smook, 1992) · Type of wood (hardwood or softwood) · Wood chip quality (size and thickness uniformity) · Wood chip moisture content · Sulfidity of cooking liquor · Strength of chemical application · Liquor-to-wood ratio · Temperature cycle · Cooking time/Reaction rate · Extent of delignification (kappa number) · Residual alkali

The kraft process can be carried out in a reacting scheme that is batch or continuous. Advantages of batch and continuous schemes are delineated in Table 8. The vessel in which the cooking takes place is commonly referred to as a digester. The three-phase behavior in the digester is characterized by the wood chips immersed in the aqueous solution, with the gas phase being steam injected for heating purposes and/ or volatile organic gases produced as the reaction proceeds. Cooking in a batch digester generally takes 2–4 hours at a maximum temperature around 170°C, see Table 9. The maximum temperature is typically reached in 1.0 to 1.5 hours, allowing the cooking chemicals to penetrate the wood chips. The cooking reactions then proceed for up to 2 hours at the maximum temperature. The heating of the cooking mixture is conducted through direct steam injection or by forced circulation of the cooking chemicals through an external heat exchanger. Batch digesters are fabricated from carbon steel with a size range of 200 to 265 m3 and are capable of producing up to 19 tons of pulp per batch (Smook, 1992).

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Table 8. Advantages of Batch and Continuous Digester Schemes (After Smook, 1992) Batch Digester Advantages

Continuous Digester Advantages

· More reliable operation

· Lower steam requirement (less energy)

· Greater operating flexibility

· More constant steam demand

· Ability to change type of pulp produced

· More compact; less space requirement

· Ability to cook softwood and hardwood concurrently

· Lower capacity requirement of auxiliary components due to constant loading

· Ease of start-up and shut-down

· Adaptable for digestion of all types of wood

Table 9. Cooking Conditions for the Kraft Chemical Pulping Process Raw Material:

Wood Chips

Chemicals:

NaOH, Na2S anthraquinone, polysulfide (extended)

Cooking Time:

2-4 hours

Cooking Liquor pH:

≈ 13

Cooking Temperature:

170-180 m C

Yield:

40-50%

The increasing demands on the pulp and paper industry to reduce emissions from bleaching operations has lead to modifications of the cooking process to provide a pulp that requires a decreased amount of bleaching. This is accomplished by decreasing the amount of residual

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608 Fluidization, Solids Handling, and Processing lignin content in the unbleached pulp by extended delignification in the cooking process. The most widely researched method to achieve extended delignification is the addition of additives, such as surfactants, anthraquinone, or polysulfide, to the cooking process. The additives work by either accelerating the delignification or enhancing the penetration and diffusion of the cooking chemicals into the wood chips (Chen, 1994). An example of a mechanism for improved penetration and diffusion through the addition of a surfactant was provided by Chen (1994). The surfactant, Chen (1994) stated, assisted in the removal of hydrophobic extractables that block the cooking chemicals from penetrating the chips. The uniformity of the pulp was also increased under these conditions. A desirable feature of extended delignification is the minimal capital cost requirement for digester modification or retrofitting (McDonough, 1995). A requirement that has received attention in the literature is a provision for improved removal of the pulp from the digester at the completion of cooking. Traditionally, the pulp has been removed from the digester by blowing the contents under full digester pressure into an adjacent blow tank. This type of removal subjects the pulp to complex two- or three-phase flows with high velocities; the pulp experiences severe temperature and pressure drops and high velocity impacts with the blow tank walls. Therefore, pulp sampled from the digester at the end of cooking but before blowing demonstrated greater strength than pulp tested after the blowing process. Cyr et al. (1989) modified the discharge arrangement on a batch digester by adding dilution liquor and pumping out the contents of a cooled and depressurized digester at a low and controlled flow velocity. With this discharge arrangement, the strength levels of the discharged pulp approached those of the unblown material in the digester at the end of cooking. Another area where a major research effort has been undertaken is in the automated control of batch digesters. Improved controlling devices for batch digesters are implemented to improve the uniformity of the pulp, reduce production costs, and save energy (Dumont, 1986). The control of batch digesters is complicated because of the inability to measure the kappa number of the pulp in the digester on-line (Lee and Datta, 1994). During a batch cook, the extent of delignification must, therefore, be conducted by secondary measurements. Recall that the kappa number is a measure of the residual lignin content in the pulp and a batch cook

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generally targets a set kappa number to reach at the end of the cook. The targeted kappa number depends on the time, temperature, and type of wood chips (hardwood or softwood) in the digester among other factors. The attainment of the targeted kappa number in the digester is critical in decreasing the chemical load in the bleaching of the pulp and, hence, reducing the environmental impact. The secondary measurements on the extent of delignification focus on determining liquid properties (Paulonis and Krishnagopalan, 1988; Lee and Datta, 1994), such as liquid temperature, lignin concentration, solid concentration, and effective alkali, or also by estimating the extent of carbohydrate degradation during cooking (Alen et al., 1991). Lee and Datta (1994) provided a review of the various empirical models, based on secondary measurements, used to assist in the control of a pulp digester. The empirical nature of the models leads to difficulties in handling major changes in the process. Examples of control techniques include a nonlinear approach using liquid analysis (Lee and Datta, 1994), the use of neural networks (Dayal et al., 1994; Yeager, 1995) and continuous process improvement techniques (Saraiva and Stephanopoulos, 1992). Other procedures to improve and optimize pulp digester operation lie in the improvement of scheduling a set of batch digesters (Hvala et al., 1993; Rihar, 1994) and improvements in the heat economy of the digesters from batch to batch (Tikka et al., 1988; Petrov…i… et al., 1995). While a review on all the various types of control techniques is beyond the scope of this section, the highlighted control procedures/techniques are the most commonly encountered in the literature. 3.3

Pulp Bleaching and Flotation De-inking

The objective of bleaching the pulp is to continue the delignification process started in cooking, thereby removing materials which contribute to the pulp color and leaving behind cellulose and hemicellulose which are inherently white. The bleaching process consists of a sequence of process stages, applying different chemicals in each stage. The most common chemicals used in bleaching stages are listed in Table 10. Based on the bleaching chemicals (i.e., chlorine, oxygen, ozone), the bleaching of pulp occurs in a three-phase mixture. The use of elemental chlorine (Cl2) has decreased in recent years because of environmental concerns about the

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610 Fluidization, Solids Handling, and Processing linkage of chlorine bleaching and dioxin in the effluent from the bleaching plant. Elemental chlorine-free (ECF) bleaching sequences, which substitute chlorine dioxide (ClO 2) for elemental chlorine, are the primary choice presently in the paper and pulp industry. Environmental pressures to completely eliminate chlorine based products from bleaching sequences has lead the industry to explore new chemicals and processes to achieve totally chlorine-free bleaching (TCF) in the near future. The high oxidizing potential of ozone has put ozone bleaching technology at the forefront of TCF bleaching. The barriers to overcome in the application of ozone to TCF bleaching are the poor selectivity of ozone (ozone oxidizes cellulose as well as lignin), nonuniform bleaching with ozone that can adversely affect the pulp, and the high cost associated with the production of ozone (Nutt et al., 1993). Table 10. Common Chemicals used in Bleaching of Paper Pulps

Chemical

Conditions

Chlorine

Acidic medium

Chlorine Dioxide

Acidic medium

Caustic

Alkaline extraction

Oxygen

Alkaline medium

Sodium Hypochlorite

Alkaline medium

Hydrogen Peroxide

Alkaline medium

Ozone

Acidic medium

The bleaching reaction is traditionally conducted in vertical towers approximately ten meters in height with the pulp slurry mixed with the bleaching gas at the inlet and pumped continuously through the tower. The three-phase reacting mixture travels through the column, via upflow or downflow, as a moving bed or plug flow type system. The mixing of the gas and pulp slurry at the inlet, therefore, becomes critical in providing a

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uniformly bleached pulp. Coalescence of bubbles can lead to a decreased reaction rate because of a decrease in the surface area of the bubbles in contact with the pulp slurry and an increase in channeling of the gas phase through the reactor, which decreases the residence time of the gas phase in the reactor. These bleaching reactions can be carried out under a wide range of weight percent of pulp or consistency (based on the oven dry weight of the pulp in the slurry) of the pulp phase. The choice of the percentage of pulp or pulp consistency at which the bleaching will take place is important due to several factors. The first is that the bleaching reaction occurs quickly once the bleaching chemical comes into contact with the pulp phase, especially in the case of ozone. This means that the bleaching reaction is mass transfer limited. The greatest resistance to mass transfer in a well mixed system occurs in the liquid phase that separates the gas bubble from the pulp fibers; i.e., in the mass transfer from the bulk gas phase to the surface of the fiber, see Fig. 4. Therefore, a reduction in the amount of liquid present (an increase in pulp consistency) decreases the resistance to mass transfer (White et al., 1993). Increasing the pulp consistency, however, renders the pulp less amenable to mixing in traditional mixers and more complicated designs for high consistency (25% to 28%) bleaching, in some cases, are required. An example of this, shown in Fig. 5, is the high consistency ozone bleaching paddle conveyor reactor developed by Union Camp Corporation, which is a horizontal, tubular reactor with shaft-mounted paddle-type internals (White et al., 1993). High and medium (10–14%) consistency bleaching reduces the volume of pulp slurry to be pumped and the amount of effluent to be treated at the cost of increased mixing requirements and, possibly, more complicated reactor design. Hurst (1993) demonstrated that the effectiveness of ozone as a bleaching chemical increased with increasing pulp consistency under similar reaction conditions (mixing intensity, reaction time). The downside of high consistency ozone bleaching is a loss in the viscosity of the pulp (cellulose); however, Kappel et al. (1994) reported that the strength properties were not significantly affected. Hurst (1993) also stated that increasing the reaction time for low consistency (3–9%) pulps provided results similar to the high consistency case.

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20,000) via the following two-step reaction: CO + 2H 2 → (- CH 2 -) + H 2 O CO + H 2 O → H 2 + CO2

∆ H r (227°C) = -165 kJ/mol

∆ Hr (227°C) = - 39.8kJ/mol

The overall reaction of the desired hydrocarbon synthesis is thus

2 CO + H2 → (-CH2 -) + CO2

∆ Hr (227 °C) = - 204.8 kJ/mol

Selectivity to desired products including light hydrocarbons, gasoline, or diesel fuel depends upon the catalyst employed, the reactor temperature, and the type of process employed. Products of the F-T synthesis are suitable for further chemical processing because of their predominantly straight chain structure and the position of the double bond at the end of the chain. By-products formed on a lesser scale include alcohols, ketones, acids, esters, and aromatics.

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620 Fluidization, Solids Handling, and Processing The catalysts used in the F-T synthesis are produced from iron precipitation and are typically less than 50 µm in size. With the small size of the catalyst particles, the three-phase reacting system is commonly referred to as a slurry bubble column. The highly exothermic nature of the desired synthesis is the cause of two of the main technical problems associated with the F-T synthesis: that of removal of the heat of reaction, amounting to about 25% of the heat of combustion; and avoidance of local overheating of the catalyst, which favors methane formation and catalyst deactivation (Kolbel and Ralek, 1980). The use of a slurry bubble column reactor effectively eliminates these problems by virtue of its excellent heat transfer characteristics and temperature uniformity. Other advantages offered by the slurry bubble column over fixed beds or moving beds in the absence of liquid media include high catalyst efficiency; favorable catalyst regeneration conditions; flexibility in establishing operating conditions, which allows economical control of the synthesis; ease of addition or removal of catalyst; and simplicity of construction and, therefore, lower capital costs. One disadvantage of the slurry bubble column in this case is the degree of backmixing of the liquid and, to a lesser extent, the gas phase, which limits the conversion and specific product selectivity attainable. The backmixing problems, however, can be reduced by using baffles. Transport properties and reaction engineering aspects of F-T synthesis have been the subject of comprehensive reviews, e.g., Kolbel and Ralek (1980), Anderson (1984), and Saxena et al. (1986). Readers are referred to these references for details on the F-T synthesis and the extensive literature listings contained in them. One recent development in low-temperature F-T synthesis that requires mentioning is the development by Sasol of a commercial scale Slurry Phase Distillate process, which produces high quality fuels, solvents, paraffins, and waxes from hydrogen and carbon monoxide. The process is located in a plant in Sasolburg, South Africa and consists of a 4.8 m diameter slurry bed reactor capable of producing environmentally friendly diesel from natural gas, providing the opportunity to convert remote natural gas reserves into low aromatic content, sulfur-free diesel fuel that burns cleanly with low particulate emissions. Sasol believes the reactor technology and operational mode provides the capability for performing exothermic reactions in slurry phase catalytic systems beyond FT synthesis (Inga, J. R., personal communication, 1994).

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621

Another current development in the use of F-T chemistry in a threephase slurry reactor is Exxon’s Advanced Gas Conversion or AGC-21 technology (Eidt et al., 1994; Everett et al., 1995). The slurry reactor is the second stage of a three-step process to convert natural gas into a highly paraffinic water-clear hydrocarbon liquid. The AGC-21 technology, as in the Sasol process, is being developed to utilize the large reserves of natural gas that are too remote for economical transportation via pipelines. The converted liquid from the three-step process, which is free of sulfur, nitrogen, nickel, vanadium, asphaltenes, polycyclic aromatics, and salt, can be shipped in conventional oil tankers and utilized by most refineries or petrochemical facilities. The three-step process for the AGC-21 technology is shown in Fig. 7. The three-phase slurry reactor contains an internal heat transfer coil to provide effective heat removal from the exothermic reaction. The feed to the hydrocarbon (F-T) synthesis step is a hydrogen/carbon monoxide synthesis gas (syngas) with 2/1 H2/CO ratio. The product from the slurry reactor is a high wax content material, which undergoes further upgrading in the third step. The catalyst is multimetallic, cobalt-based solid particles optimized for operation in a slurry environment (Eidt et al., 1994). Separations are required to recover the product from the catalyst containing slurry. Demonstration of the process has been conducted at Exxon Research and Development Laboratories in Baton Rouge, Louisiana (Everett et al., 1995) with a slurry reactor which was 1.2 m in diameter and approximately 21 m in height. The design and development target for the AGC-21 process are facilities capable of producing 8–16 km3/day of liquid hydrocarbon (Eidt et al., 1994). 4.4

Methanol Synthesis

Synthesis gas can be reacted or reformed to make methanol through highly exothermic reactions: CO + 2 H 2 ↔ CH 3 OH (l)

∆ H r (255°C) = - 91.9 kJ/mol

CO 2 + 3 H 2 ↔ CH 3 OH (l) + H2 O

∆ H r (255°C) = -51.2 kJ/mol

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622 Fluidization, Solids Handling, and Processing The first reaction produces methanol with a low hydrogen consumption, but evolves significantly greater amounts of heat. The second reaction evolves less heat, but consumes more hydrogen and produces the byproduct steam. Thermodynamically, low temperatures and high pressures favor methanol formation. The reactions are carried out with coppercontaining catalysts with typical reactor conditions of 260°C and 5 MPa (Probstein and Hicks, 1982).

Syngas Generation CH.

+ tOZ -CO

Hydrocarbon S~nthesis + 2Hz

.-C-C-C-C-C

H):droisomerization

~

C-C-~-C

Figure 7. Exxon’s AGC-21 process for conversion of natural gas into liquid hydrocarbons. (From Eidt et al., 1994.)

Three-phase reactor systems are ideally suited for methanol production because of the ability to provide intimate contact between the gaseous phase reactants and the solid phase catalysts and to remove the large amounts of heat created by the high heats of reaction. In the three-phase system, an inert liquid phase circulates between the reactor and an external

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623

heat exchanger providing the necessary heat removal. The inherent temperature uniformity of three-phase systems increases catalyst life, insures optimum selectivity and reaction rate, and reduces NOX formation. The three-phase systems can be operated in a fluidized bed mode with catalyst particles between 3 and 6 mm in size or in a slurry bubble column mode with catalyst particles about 50 µm in size. Research efforts in catalyst activation procedures, choice of liquid medium, reactor hydrodynamics, and reactor modeling are currently being pursued for improving the process efficiency, design, and evaluations (e.g., Toseland et al., 1995).

Part II: Three-Phase Biofluidization 5.0

BIOLOGICAL APPLICATIONS OF THREE-PHASE FLUIDIZATION

5.1

Introduction

The many positive features of three-phase fluidization are readily apparent to biochemical engineers. Improved mass transfer, improved solid and liquid mixing, elimination of clogging, and other advantages of the technique as compared to packed bed and suspended cell bioreactors, detailed in Table 11, have been confirmed for a wide spectrum of biological processes, such as wastewater treatment, antibiotic production, and animal cell culture. By immobilizing cells on or in particles, washout of biomass is prevented even at relatively high flow rates, unlike the situation in continuous suspended cell reactors. Several limitations to the technique, listed in Table 12 exist; however, the favorable aspects of threephase biofluidization often outweigh the negatives and have encouraged development of many applications, especially in wastewater treatment. Numerous examples of full-scale three-phase fluidized wastewater treatment facilities exist (Fan, 1989). Industrial fermentation and cell culture applications have been somewhat slower to develop, but are becoming more and more common.

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624 Fluidization, Solids Handling, and Processing Table 11. Advantages and Positive Features of Three-Phase Fluidized Bed Bioreactors Advantages and Positive Features immobilized cells

• higher cell density results in higher productivity • no washout at high flow rates

References Andrews and Przezdziecki, 1986; Kargi and Toprak, 1994

• may eliminate need for cell recycle greater stability

• tolerates periods of downtime • less affected by shocks of toxic material (no washout, adsorbent particles)

Livingston and Chase, 1991; Yongming et al., 1993

degree of axial liquid mixing can be varied

• can tailor mixing to biokinetics

Andrews and Przezdziecki, 1986; Fan, 1989

freely fluidized particles

• no clogging and easy dispersal of gas evolved during fermentation

Andrews and Przezdziecki, 1986; Qureshi and Maddox, 1988; Davison and Scott, 1992; Lazarova et al., 1994

• possibility to add/remove fourth phase for product removal • can use smaller particles than packed beds, resulting in improved intraparticle mass transfer • low pressure drop operational flexibility

• can treat high strength wastewaters at low rate or dilute wastewater at high rate

Borja and Banks, 1994

mechanical simplicity

• simple construction sterility

Siegel and Robinson, 1992; Kargi and Toprak, 1994; Merchuk et al., 1994

• more easily maintained • relatively low capital cost

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625

Table 12. Limitations of Three-Phase Fluidized Bed Bioreactors Disadvantages

References

insufficient data and/or models available

scale up difficult sparse data from actual fermentations data

Davison, 1989; Schiiger1, 1989; Siege! and Robinson, 1992

potential for undesired particle stratification

affects substrate concentration to which biomass is exposed

Andrews and Przezdziecki, 1986; Lele and Joshi, 1992

aeration rate affects

recycle rate can be varied to affect both

Merchuk et al., 1994

lower potential volumetric productivity

Shu and Yang, 1996

agitation

potentially lower cell density than in packed beds

Though the various biological applications of three-phase fluidization are at somewhat different stages of development, these areas share similar process considerations that are not relevant to three-phase fluidized chemical processes, primarily as the result of the complex interaction between the biological systems and the reactor hydrodynamics, mixing characteristics, and transport phenomena. This interaction is schematically depicted in Fig. 8. The biological reaction is the goal of the process, thus the effect of the reactor characteristics on the biokinetics and vice versa are critical to the design and operation of a three-phase fluidized bed bioreactor. This section of the current chapter goes beyond a simple listing of current three-phase biofluidization applications to consider the differences in conventional three-phase fluidization and biofluidization from the aspect of reactor design and operation. Past research into three-phase biofluidization has been summarized in several excellent reviews (Andrews, 1988; Fan, 1989; Heijnen et al., 1989; Schügerl, 1989; Siegel and Robinson, 1992), and this chapter will concentrate on the main research themes and advances of the last few years. Though gas continuous three-phase fluidized bioreactors exist (Fan, 1989), we consider here only those bioreactors in which the liquid phase is the continuous phase.

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Figure 8. The relationship between biological, hydrodynamic, mixing, and transport phenomena in three-phase fluidized bed bioreactors.

The definition of phases in a three-phase fluidized bioreactor is not always as clear as in traditional three-phase fluidization applications. The solid phase is made up of bacterial, fungal, plant, or animal cells, immobilized on or in some type of support or self-immobilized in the form of flocs or pellets. If a support particle is not present, the reactor is considered to be a three-phase bioreactor if the bioflocs or cell aggregates are greater than 10 µm, at which scale they can be considered heterogeneous to the liquid phase in terms of fluid dynamics (Fan, 1989). The liquid phase constitutes the medium for cell growth. The gas phase may be supplied to the reactor through a distribution plate, as is typical for other three-phase fluidization applications, or the gas phase may be generated within the reactor as part of the biochemical reactions taking place, or both. Depending on whether the fermentation is aerobic or anaerobic, the gas phase may be air, oxygen, or an inert gas. If evolved during the fermentation, the gas phase may be, for example, carbon dioxide, methane, or hydrogen. There are a wide variety of three-phase fluidized bioreactor designs possible. The conventional reactor, shown in Fig. 9, is fluidized by both gas and liquid entering at the bottom of the reactor and leaving at the top and is the most common type of three-phase fluidized bed bioreactor. This reactor may be configured to operate with little axial liquid mixing or in a well-mixed mode by adding a recycle stream. The airlift reactor or draft tube fluidized bed reactor, Fig. 10, is also frequently used. In this reactor, gas is injected at the bottom of a draft tube placed in the center of the

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627

column, and a cyclic flow of liquid, particles, and, sometimes, gas results. The draft tube fluidized bed bioreactor is well-mixed and has excellent mass transfer characteristics (Kang et al., 1991). It is well understood in bioprocessing that process parameters, such as temperature, substrate concentration, and pH, affect bioreactor performance (Nagamune et al., 1988), but certain parameters not typically considered in bioreactor design, such as particle size or concentration and biofilm thickness and morphology, can dramatically affect the stability of a three-phase fluidized bed bioreactor (Oh et al., 1988). Table 13 lists many of the parameters that must be considered by the biochemical engineer in order to successfully design and operate such a bioreactor. Many of the system parameters are discussed in detail by Fan (1989) and will not be discussed further here. A comprehensive discussion of the biological parameters is beyond the scope of this chapter, and the reader is referred to the abundant literature in this field; however, this section of the current chapter attempts to provide an overview of the key biological considerations for the design and operation of three-phase fluidized bioreactors. G

.0 .

6. .0

G

Liquid Gas

Stream Stream

ProdlXt

L .I o

mmobilized Gas

Cell

Part"K:les

Bubbles

Stream

Figure 9. Conventional three-phase fluidized bed bioreactor.

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L: Liquid

Stream.

G: Gas Stream P: Prod.r:t Airlift

I mmobilized o Gas

Cell Partcles

Bubbles

Stream

or draft-tube three-phase fluidized

bed bioreactor.

System and Biological Parameters that Affect Bed Bioreactor Performance

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Three-Phase


629

Applications

The most widespread biological application of three-phase fluidization at a commercial scale is in wastewater treatment. Several large scale applications exist for fermentation processes, as well, and, recently, applications in cell culture have been developed. Each of these areas have particular features that make three-phase fluidization particularly well-suited for them. Wastewater Treatment. As can be seen in Tables 14a to 14d, numerous examples of the application of three-phase fluidization to wastewater treatment exist. Laboratory studies in the 1970’s were followed by large scale commercial units in the early 1980’s, with aerobic applications preceding anaerobic systems (Heijnen et al., 1989). The technique is well accepted as a viable tool for wastewater treatment for municipal sewage, food process waste streams, and other industrial effluents. Though pure cultures known to degrade a particular waste component are occasionally used (Sreekrishnan et al., 1991; Austermann-Haun et al., 1994; Lazarova et al., 1994), most applications use a mixed culture enriched from a similar waste stream or treatment facility or no inoculation at all (Sanz and FdezPolanco, 1990). Table 14a. Recent Three-Phase Biofluidization Research on Treatment of Municipal Wastewater

AC = activated carbon

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630 Fluidization, Solids Handling, and Processing Table 14b. Recent Three-Phase Biofluidization Research on Treatment of Toxic Wastewater

I

~~~~

I

Stream phenol

OrganIsm

Mixed culture

Carrier

Bead

Comments

Rererences

sand

draft tube improved nuidizalion; stagedreacto"gaveexcellenl

activated carbon

drafl tube nuidized bed

TzenB, 1991

draft lube conlrolled biofilm Ihickn...

Fan el al.. 1987b

FBR was less affected than conventional systems by low

Tseng and Lin, 1994

!

Etzensperger et al., 1999

degradation

temperalure anlhracite coal

oscillations of phenol concentntion in emuent in response to inlel change

Worden and Donaldson,1987

activated carbon

biokinelic paramele" and phenol diffusivity in biofilm estimated

Tang and Fan, 1987

Pseudomo""s sp.

PVA.alBinate beads

counter-cumnl wilb aeration

Wu and Wisccarvcr. 1989

nonacclimated activated sludge

Celite1M, a silica. based support

low excess biomass production

M~ncnctal,I993

sand; Celile R. 6331M; volcanite;

high rate bio~medic;ation allow t.mperatu~s (4-10 "C) dcmonslraled

lKrvinen.1 aI., 1994

unspecified unspecified

Pseudomonas spp. dominale biofilm after two years maintenance in continuous aerobic nuidized bed ~aclor

Puhakka Puhilla et el at.. 19 1995

biomass capable of degrading ~calcitrant compound was slow growing and ~quired special stArt-up procedu~ -packed bed biofilm

Livingston.

sewage from coal conve"ion plant

chlorophenols

nuidized bed reactor

pumice

dichloroaniline

commercial starter culture

diatomaceous earth

polluled soil

mixed soil organisms

sand

tapered column allows stratification fine and COIr'" particles; independently adjustable ~tention times

sand

2-stage system of recycled sludgeblankel ~aclor followed by nuidized

1991

development

nilrlle and selenium

Tha".ra S./.Mti.t

bed reaclor

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of

van der lagt et at.. 1993

Macy el at.. 1993


631

Table 14c. Recent Three-Phase Biofluidization Research on Treatment of Food Plant Wastewater

Table 14d. Recent Three-Phase Biofluidization Research on Treatment of Industrial Wastewater

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632 Fluidization, Solids Handling, and Processing The advantages of three-phase fluidization are especially suited to wastewater treatment. For instance, the high biomass retention inherent in immobilized cell reactors allows high pollutant loading, and the ability of the reactor to handle suspended solids without clogging is ideal for many sewage streams (Andrews, 1988; Weuster et al., 1989). Three-phase fluidized bed bioreactors allow for flexible operations; flow rate and pollutant loading can be varied over a wide range, a requirement in many sewage treatment systems (Tzeng, 1991). The high productivity of threephase fluidized bed bioreactors results in relatively compact units. This is of particular value when a given substrate can not be degraded by the wastewater treatment bioreactor that treats general plant effluents; a threephase fluidized bed bioreactor with an inoculum tailored to the recalcitrant substrate can easily be placed at the point of generation of that substrate, treating it before the waste streams are commingled. Current work in this area is focusing on improved carrier materials, tailoring the reactor/inoculum system to a specific waste stream, improving reactor start-up, and determining optimum operating conditions. Low temperature treatment has been shown to be feasible (Sanz and FdezPolanco, 1990), and responses to shock loading of organics have been elucidated in several cases (Zellner et al., 1994; Borja and Banks, 1995). Adsorption and desorption of degradable substrates by the support particle can be effective in preventing toxic shock caused by sudden increases in the substrate concentration and can prevent excess amounts of the substrate in the reactor effluent. The ability of the support particle to desorb the substrate is critical in allowing regeneration of the particles (Mol et al., 1993; Petrozzi et al., 1993). Automatic control of pH has also been studied as a method of protecting the reactor from toxic surges in effluent pollutant concentrations, improving reactor stability and effluent quality, and reducing costs for alkali to treat reactor imbalances (Denac et al., 1990; Ehlinger et al., 1994). Conditions during start-up can have a dramatic impact on time required to reach reasonable performance levels and on the ultimate reactor performance. Dilution rate, loading rate as measured by chemical oxygen demand (COD), carrier choice, inoculum amount, inoculum strategy, and inoculum species distribution are critical parameters (Sreekrishnan et al., 1991; Araki and Harada, 1994; Austermann-Haun et al., 1994; Yongming et al., 1993). Pore characteristics have a strong influence on time required for start-up and on the ultimate biomass density in the

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633

reactor; high porosity, adequate pore size, and large surface area all contribute to rapid reactor start-up and high cell density, though no single one of these parameters by itself is an adequate predictor of optimal cell retention for successful reactor start-up and operation (Yee et al., 1992). Control of reactor pH by controlling the feed rate has been shown to result in more rapid reactor start-up than without pH control (Hsu and Shieh, 1993). During start-up, the microbial population distribution in the biofilm varies with time. Initial colonization of the particle may be by one or more species that alter the surface favorably for colonization by other species. For instance, in the operation of a butyrate-degrading fluidized bed bioreactor, methanogens attached to the sand particles early in the start-up process and produced a primary matrix of heteropolysaccharides that allowed attachment of other bacterial species (Sreekrishnan et al., 1991; Zellner et al., 1991; Yongming et al., 1993). This is contrary to findings in an acetate-propionate-butyrate degrading reactor, in which facultative anaerobes were found to be the initial colonizers (Lauwers et al., 1990). The COD loading profiles used for start-up may be based on maximum efficiency or maximum loading. In the former, COD loading is gradually increased and residence time is gradually decreased as the biofilm develops; high COD removal percentages are achieved early in the reactor lifetime, but initial loading is low. The maximum loading profile starts with a short residence time and high COD loading, resulting in initially low COD removal efficiencies; the initially high substrate concentrations have been suggested to increase the rate of reactor start-up (Hsu and Shieh, 1993; Yongming et al., 1993). One of the problems associated with waste treatment is the presence of recalcitrant or nonbiodegradable compounds in the waste stream (Keller, G., personal communication, Nov. 3, 1995). If the bioreactor cannot degrade toxic compounds that the law requires to be removed, either an alternative chemical or physical process must be completely substituted for the bioprocess or a supplemental process must be added to the treatment system. The combination of biodegradation with physical adsorption in a fluidized bed reactor has been successfully demonstrated for the treatment of cyanide-containing, phenolic wastewaters (Pfeffer and Suidan, 1989); the biofilm was immobilized on granulated activated carbon which acted as a sink for both the degradable substrate and toxic, nondegradable compounds. Fresh carbon was periodically added to the reactor, replacing exhausted carbon, so that the toxic nonbiodegradable compounds did not

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634 Fluidization, Solids Handling, and Processing build up in the bulk reactor fluid. Because much of the degradable substrate was consumed by the biofilm, carbon requirements were reduced from what they would be if no biodegradation took place (Pfeffer and Suidan, 1989). Fermentations. Recent years have seen many investigations of fermentations to produce high-value products, such as pharmaceuticals and recombinant proteins, in fluidized bed bioreactors. Continuous production of these products in three-phase fluidized bed bioreactors is especially attractive, because the simplicity of construction of these bioreactors reduces the opportunities for contaminating organisms to gain entry to the reactor during long-term production. The advantages of threephase fluidized bed bioreactors have also been applied to more traditional bacterial fermentations, such as lactic acid or solvent production. For fermentations such as these that exhibit product inhibition, the potential to integrate product removal with the fermentation by adding adsorbent particles as part of the solid phase of the fluidized bed is a unique advantage currently being explored by several researchers (van der Wielen et al., 1990; Davison and Scott, 1992; Kaufman et al., 1995). Though several commercial scale three-phase fluidized bed bioprocess operations exist (Fan, 1989), much of the work in applying three-phase fluidization to fermentation is still at a laboratory scale. A partial survey of this burgeoning research is given in Table 15. Fermentation applications in three-phase fluidization have lagged those of wastewater treatment, because the goal, biosynthesis of useful chemicals, is more constrained than that of simple biodegradation of waste. Sterilization, pure culture preparation, separation of intermediate or product inhibitors, bioreactor control, specialized techniques of microbial immobilization, and product recovery must be considered. Oxygen demands are generally higher for aerobic fermentations than they are for aerobic wastewater treatment applications, because the higher substrate concentrations in a fermentation feedstock result in higher metabolic requirements for oxygen (Andrews, 1988). Several areas are receiving much of the research attention. Approaches that integrate product recovery with the fermentation in a threephase fluidized bed bioreactor reflect general research trends in biochemical engineering (Yabannavar and Wang, 1991; Davison and Thompson, 1992). The successful use of three-phase biofluidization has also been demonstrated for recombinant protein systems, where it may have some benefit in improving plasmid stability (Shu and Yang, 1996).

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635

Table 15. Recent Applications of Three-Phase Biofluidization to Fermentation Processes

Cell Culture Applications. Animal cell immobilization on microcarriers is a widely used technique in animal cell culture (Kratje et al., 1994) and is as easily applied to biofluidization as to the stirred tank applications for which it was initially developed. Cell culture applications, which frequently require long culture times, benefit from the mechanical simplicity of the three-phase fluidized reactor because of the reduced likelihood of contamination. Avoidance of high intensity shear caused by agitation (Venkat et al., 1996) is another advantage of the use of

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636 Fluidization, Solids Handling, and Processing three-phase fluidized bed bioreactors for cell culture, especially for bioreactors of the airlift or draft tube type (Siegel and Robinson, 1992). Table 16 lists a number of recent cell culture applications for three-phase fluidization. These include production of monoclonal antibodies and therapeutic proteins, such as erythropoietin. Plant cells have also been cultured in airlift reactors (Scragg, 1992). The combination of three-phase fluidization and cell culture may represent an ideal union of new technology and immediate application; because the high-value products of mammalian cell culture require relatively small production scales (on the order of 100 liters rather than 100,000 liters), it is possible to investigate the usefulness of three-phase fluidization in what would otherwise be considered a pilot scale set up without having to build an expensive, large scale unit. Table 16. Recent Applications of Three-Phase Biofluidization to Cell Culture

BHK = baby hamster kidney cells SS = stainless steel CHO = chinese hamster ovary cells

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637

Bioparticles

The type of bioparticle chosen for three-phase biofluidization often differs from that used in traditional three-phase fluidization. Small, dense particles, similar to those used in traditional catalytic three-phase fluidization applications, were often used in the initial studies of biofluidization, and for some time it was generally assumed that a bioparticle consisted of a biofilm grown on a small, dense particle (Heijnen et al., 1989). There are presently many applications, however, using cells entrapped in low density gel beads or growing within low density porous particles (Fan, 1989). These low density particles (alginate beads, 1.05 g/cm3; biofilm covered sand, 1.07 g/cm3) cause striking differences in the hydrodynamic characteristics of the biological systems compared to particles traditionally used in three-phase fluidization, such as glass beads (2.5 g/cm3). Furthermore, the density of the particle often changes with time during the bioreaction because of cell growth, again affecting the hydrodynamics of three-phase fluidized bioreactors. Immobilization Methods and Particle Selection. Many methods of immobilization have been developed; these and their advantages and disadvantages have been well reviewed (Fan, 1989) and will not be discussed in detail here. Immobilization by the natural formation of a biofilm on a solid support and by the entrapment of cells within a polymer such as calcium alginate are the methods most frequently used in biofluidization research. The growth of animal cells on and in microcarriers, such as macroporous sintered glass spheres, is also a widely studied method of immobilization for three-phase biofluidization. Choice of method and particle material are determined by effects of particle characteristics on reactor performance. Desirable particle qualities include ease of handling and preparation for immobilization, reusability, suitability for steam sterilization, and low cost (Gonçalves et al., 1992). Table 17 lists several examples of particles in use in three-phase biofluidization research. It must be recalled in selecting a particle that cell immobilization of any means can affect biological processes through various physicochemical effects, such as partitioning and diffusivity, and by directly affecting cell physiology and morphology. For instance, the optimal process pH for immobilized cell systems is likely to differ from the optimal pH for suspended cell culture because of the effects of surface phenomena on local pH values (Gonçalves et al., 1992).

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638 Fluidization, Solids Handling, and Processing Table 17. Examples of Particle Materials used for Various Three-Phase Fluidized Bioreactor Applications

The particle materials also affect the ability of various organisms to colonize the particle and form an active biofilm through complicated mechanisms of surface charge, roughness, and porosity (Fan, 1989; Grishin and Tuovinen, 1989; Gonçalves et al., 1992). Cells growing within particle pores may be protected from shear effects (Grishin and Tuovinen, 1989; Livingston and Chase, 1991). Colonization ability has been shown to vary with species as well as with particle material (Lazarova et al., 1994). Even in studies where biomass holdup was constant, the support material affected the biomass activity (Ruggeri et al., 1994). Attrition of particles can also be important; because the attrition rate varies with superficial fluid velocity (Nelson and Skaates, 1988), a candidate material must be tested under the range of conditions anticipated in the fermentation. Colonization is also affected by process conditions such as fluid velocities and even influent composition, underscoring the need for particle screening under the expected process conditions (Gonçalves et al., 1992; Mol et al., 1993).

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639

Biofilm Effects. Particle size and density are especially important in determining the success of biofluidization. Bioparticle size and density are determined by the initial particle properties, the extent of biomass growth, and biofilm density. For a biofilm covered support, the apparent particle density is determined by Eq. (7) (Fan, 1989). 3

Eq. (7)

d  ρ p = ( ρ s - ρ bw )  o  + ρ bw  dp   

Particle sizes used for aerobic waste treatment generally range from 100– 1000 µm in size, with most being less than 500 µm in diameter, and reported biofilm thicknesses range from 40–1200 µm, with 100–200 µm being typical (Fan, 1989). The terminal settling velocity of the particle is determined by the particle size and density as shown in Eq. (8) (Kunii and Levenspiel, 1991):

Eq. (8)

 4 d p ( ρ - ρ )g  s f  ut =    3 ρ f CD  

0 .5

The effect of biofilm volume on the terminal settling velocity can be described by Eq. (9) (Andrews and Przezdziecki, 1986): u t (1 + Bx) ( 2 - e ) = u ts (1 + x) 1 3 1

Eq. (9)

This equation is plotted in Fig. 11, showing that for relatively dense support particles, biofilm growth can reduce the settling velocity if the biofilm density is less than that of the biofilm-free particle. As such bioparticles gain biomass, they will rise to the top of the bed and may even elutriate from the reactor (Sreekrishnan et al., 1991; Myška and Švec, 1994), reducing achievable conversion rates. This situation could be resolved by using lower density particles, such as expanded polystyrene or

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640 Fluidization, Solids Handling, and Processing polyethylene beads currently used for inverse airlift reactors (Nikolov and Karamanev, 1987; Garnier et al., 1990), for which the settling velocity actually increases as the biofilm grows (Andrews and Przezdziecki, 1986).

~ u..

Figure 11. Effect of biofilm growth on terminal settling velocity.

Biofilm thickness and, therefore, bioparticle size, depends on the sum of biofilm growth and biofilm detachment. Detachment of biomass is a function of many parameters, including reactor operating conditions, particle material parameters such as surface roughness, and species specific parameters such as biofilm morphology (Chang et al., 1991). Increased medium concentration and fluid turbulence, as measured by the reactor Reynolds number, increased the rate of biofilm detachment in one study (Chang et al., 1991). This is supported by studies that showed that the removal of biomass could be controlled to a large extent by the gas and liquid velocities; increases in either increased the rate of biofilm removal (Roustan et al., 1993; Trinet et al., 1991). In cases where increasing the gas and liquid velocities is either ineffective in removing the desired

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641

amount of biomass from the particles or is undesirable from an operations standpoint, the addition of a draft tube to the reactor may provide the level of turbulence required (Livingston and Chase, 1991). Alternatively, the optimal biofilm thickness can be maintained by removing particles as the biofilm grows thick and stripping or washing some of the biofilm from the particles (Andrews and Przezdziecki, 1986; Tzeng, 1991). Modeling studies have shown that when the bed is not monosized and the difference in particle sizes is great enough, it is possible that the large particles will not reach a lower settling velocity than the small particles; hence the large particles will remain in the bottom of the bed (Myška and Švec, 1994). If biofilm thickness control is desired, removal of particles at a location other than the free surface would have to be implemented in this case. Novel Bioparticle Research. Two major thrusts have been seen in recent particle research—the area of density manipulation so that particle density suits the desired fluidization mode, and the development of magnetic particles for use in magnetically stabilized fluidization. Intraparticle mass transfer is also of interest. Table 18 lists several novel particles developed in recent years to address these and other concerns. Several methods have been suggested for increasing bioparticle density. Stainless steel mesh has been used as a support framework and density modifier for bioflocs for wastewater treatment. The large void volume available in the mesh, the capability of the relatively denser particles to withstand elutriation at high fluid velocities, thereby allowing higher reactor throughput, and the ability of filamentous organisms to be retained within the mesh were some of the advantages cited (Kargi and Toprak, 1994). Entrapment of stainless steel beads in collagen particles upon which mouse hybridoma cells were immobilized (Shirai et al., 1994) and the inclusion of α-alumina in calcium alginate beads (Paz et al., 1993) have also been demonstrated as methods to increase bead density. Though most particle density manipulation research has focused on increasing density, one report addresses the possibility of decreasing particle density to allow low aspect bioreactors with low liquid velocities required for fluidization of low density bioparticles. Hollow glass spheres were co-immobilized with yeast cells in calcium alginate beads. The final particle density after cell growth could be varied in proportion to the percentage of hollow glass spheres incorporated (Vorlop et al., 1993).

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642 Fluidization, Solids Handling, and Processing Table 18. Examples of Recently Devised Novel Bioparticles

Large, light, granular particles that allow biomass growth without alteration of particle physical properties, hence, no change in fluidization characteristics, have been used for wastewater treatment in a three-phase fluidized bed; the particle material was not specified (Roustan et al., 1993). Another approach to avoiding undesired solid stratification is to periodically remove excess biofilm from the particles (Tzeng, 1991; Livingston and Chase, 1991). Magnetic beads or support particles are used in magnetically stabilized fluidized beds. The advantages of such a bioreactor are discussed in a Sec. 5.10. A method for manufacturing a magnetic pellicular bead by electrostatically depositing alternating layers of silica and polymer onto magnetic core particles, followed by burning of the polymer and sintering of the silica, was developed to allow exploitation of the multiply derivatizable silica for various biological immobilization processes (Goetz et al., 1991). The resulting particle was dense and, thus, had increased settling velocity, allowing increased flexibility in operating conditions for the magnetically stabilized fluidized bed. This particle has been used to immobilize enzymes, but its porosity and material of manufacture suggest that it may also be suitable for microbial immobilization.

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643

Improvement of intraparticle mass transfer is the goal of some particle research efforts. One novel approach that has been recently tested is the co-immobilization of algae with bacteria; the algae produced oxygen and the bacteria produced the desired product (Chevalier and de la Noüe, 1988). Another method used microporous particles entrapped within alginate bead bioparticles to prevent excess biomass growth that could hinder intraparticle mass transfer (Seki et al., 1993). 5.4

Hydrodynamics

Predictions of bioreactor hydrodynamics and individual phase loadings are important to allow the design engineer to specify column dimensions and to size the required pumps and blowers. The hydrodynamics of three-phase fluidized bed reactors have been studied and reviewed (Muroyama and Fan, 1985; Fan, 1989), though additional work is needed to fully understand these complex systems. The necessity for further study is even greater for three-phase fluidized bed bioreactors, the hydrodynamics of which vary greatly from traditional catalytic systems (Davison, 1989; Tang and Fan, 1990; Karamanev et al., 1992; Nore et al., 1992). As described in Sec. 5.3, the low density bioparticles common in three-phase biofluidization have dramatically different hydrodynamic characteristics than do the dense particles traditionally used in three-phase fluidization. Furthermore, the fluid flow rates used in biofluidization are relatively low compared to those found in traditional fluidization, corresponding to relatively slow biological reaction rates and the need to prevent excess erosion of the biofilm or bioparticle (Tang and Fan, 1990). The hydrodynamic characteristics of a three-phase fluidized bed bioreactor depend on a number of factors, including particle properties, such as size, density, wettability, and roughness; fluid properties, such as flow rates, surface tension, viscosity, presence of surfactants, and electrolyte concentration; and reactor design, including column geometry and gas distributor design. Major differences between the hydrodynamics of low density systems, as typified by many biological processes, and high density systems include the presence of significant axial variation of individual phase holdups in the low density systems that are not apparent in high density particle systems (Tang and Fan, 1989; Tang and Fan, 1990), greater axial dispersion of solids (Bly and Worden, 1990), changes in bubble coalescence characteristics caused by the presence of fermenta-

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644 Fluidization, Solids Handling, and Processing tion medium components or by compounds produced by the growing cells (Sun and Furusaki, 1988; Bly and Worden, 1990; Béjar et al., 1992), and change in gas holdup because of biological gas production within the bed (Davison, 1989). These and other differences are further discussed in the following sections, with a view towards assisting the biochemical engineer to avoid pitfalls in designing three-phase fluidized bioreactors. Unless otherwise noted, the hydrodynamics discussed are those of conventional three-phase fluidized bed reactors, as shown in Fig. 9. Fluidization Regime. As for traditional fluidization applications, the fluidization regime—dispersed bubble, coalesced bubble, or slugging—in which a three-phase fluidized bioreactor operates depends strongly on the system parameters and operating conditions. Generally, desirable fluidization is considered to be characterized by stable operation with uniform phase holdups, typical of the dispersed bubble regime. It would be useful to be able to predict what conditions will produce such behavior. One approach recently suggested is the mapping of satisfactory and unsatisfactory biofluidization regions according to two parameters based on easily measured physical properties of the phases, such as viscosity, density, and surface tension (Béjar et al., 1992). The success of these parameters in predicting satisfactory fluidization has been demonstrated in both a straight, well-mixed, gas-fluidized bed of calcium alginateimmobilized yeast (Béjar et al., 1992; Vorlop et al., 1993) and in a tapered single-pass liquid-fluidized bed of calcium alginate-immobilized bacteria (Davison et al., 1994). Roustan et al. (1993) also applied mapping parameters to high and low density particles to predict fluidization regime with some success; however, separate maps were required for each type of particle because particle density was not included in the parameters. Further research is required to confirm the broad applicability of this approach to the prediction of whether successful fluidization is possible. Solids Holdup. The active biomass concentration in the threephase fluidized bed bioreactor is determined by the biofilm characteristics and the solids holdup. Because metabolic rates are moderate relative to chemical catalytic rates, solids holdup in biofluidization applications are generally larger than in chemical applications. Reported overall solids holdups in various three-phase fluidized bioprocesses range from 0.02 to 0.6, with most studies in the range of 0.05 to 0.4 (Sun and Furusaki, 1988; Davison, 1989; Kobayashi et al., 1990; Karamanev et al., 1992; Potthoff and Bohnet, 1993; Roustan et al., 1993; Badot et al., 1994). Solids holdup

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645

below 0.55 has been recommended to prevent particle agglomeration (Andrews, 1988). Depending on bioparticle size and density, the liquid velocity will have to be adjusted to produce suitable bed porosity. Estimation of the solids holdup for expected operating conditions, along with required solids loading, is a requirement for sizing of the reactor. The Richardson-Zaki correlation (Richardson and Zaki, 1954) 1

Eq. (10)

 n ε =  ul     ui 

is well-accepted as a description of bed expansion in liquid-solid systems, and, with some modifications, is applicable to two-phase biofluidization (Myška and Švec, 1994). Bed expansion behavior in three-phase fluidization is more complex. This is illustrated by the observed initial contraction of some beds of small or relatively low density particles (especially pertinent to biofluidization) upon the addition of gas flow (Davison, 1989; Fan, 1989; Tang and Fan, 1989; Roustan et al., 1993). Fan (1989) provides an extensive listing of expansion correlations for three-phase fluidization and describes various approaches to predicting bed expansion in three-phase fluidization. The presence of biofilm, however, may cause these correlations to be in error, and additional work is needed in this area. Gas Holdup. Determination and prediction of gas holdup is important in biofluidization because it provides an estimate of bed expansion and freeboard requirements and void space in the reactor, which is unavailable for bioreaction (Charles and Wilson, 1994; Bajpai et al., 1990). Estimates of gas holdup and expected bubble size can also assist in predicting oxygen transfer, a critical factor in aerobic bioprocesses, by providing an estimate of interfacial area for gas-liquid mass transport. Reported values for gas holdup range from 0.01 to as high as 0.45; values of 0.02 to 0.20 are typical (Davison, 1989; Tang and Fan, 1989; Bly and Worden, 1990; Tang and Fan, 1990; Potthoff and Bohnet, 1993; Sajc et al., 1995). Phase mixing is also affected by gas holdup, as described in Sec. 5.5. Many factors affect gas holdup in three-phase fluidized systems, including bead size and density, liquid physical properties, temperature, sparger type, and fluid superficial velocities (Bly and Worden, 1990). System parameters such as reactor and gas distributor design can have

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646 Fluidization, Solids Handling, and Processing such an impact on the bubble and liquid flow behavior that not only are phase holdups affected, but different bubble flow regimes can be established in two systems operated under the same conditions (Bly and Worden, 1990; Tang and Fan, 1990). Results from different bioreactors must be compared with caution; however, some generalities can be made. Overall gas holdup increases with gas velocity in the dispersed bubble regime for both low and high density particle systems (Davison, 1989; Tang and Fan, 1989; Bly and Worden, 1990; Nore et al., 1992; Potthoff and Bohnet, 1993). As gas velocity increases and the system enters the coalesced and slugging regimes, the rate of increase in the overall gas holdup decreases (Bly and Worden, 1990). The type and loading of solid particles can affect gas holdup, though the effect is somewhat unclear, undoubtedly because of system differences. Bly and Worden (1990) found that the addition of low density particles to two-phase systems generally increased gas holdups, while the addition of high density particles decreased gas holdups. This concurs with the observation of Davison (1989) that gas holdup decreased at the top of a three-phase low density particle system in the region where it became a two-phase system (above the top of the bed). Tang and Fan (1989, 1990), also working with low density particles, however, found that gas holdups decreased with the addition of the solid phase and that as the terminal velocity of the particles increased from 0.026 m/s to 0.055 m/s, the overall gas holdup decreased further. This was attributed to the higher tendency for bubble coalescence in beds of larger terminal velocity particles. Potthoff and Bohnet (1993) found no significant influence of low density solids loading on gas holdup. Karamanev et al. (1992), working with low density polyurethane particles, found that addition of solids decreased gas holdup and that the effect was greater in a draft tube fluidized bed bioreactor. In a study of the effect of electrolyte concentration on gas holdup, Bly and Worden (1990) found a strong effect. A salt solution resulted in twice the gas holdup that distilled water did under otherwise identical operating conditions, because the salt solution suppressed bubble coalescence. Investigation of this phenomenon is important in biofluidization, because biological media commonly have high electrolyte concentrations. Axial variation of gas holdup depends on operating conditions and the particular bioprocess occurring in the reactor. For a gas-producing fermentation, gas holdup increased from 0.0 to almost 0.5 with axial

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position as gas was produced until a point near the top of the bed, where the coalesced gas freely escaped as it rose into the solid-free region (Davison, 1989). In a study of three-phase fluidization of low density particles without gas-generating biomass, Tang and Fan (1989) found that the axial gas holdup variation could be regarded as negligible. 5.5

Phase Mixing in a Three-Phase Reactor

The degree of mixing of the phases within a three-phase fluidized bed bioreactor is important to reactor performance. Depending on the degree of mixing, individual bioparticles may experience different substrate and product concentrations, thereby influencing the overall reactor kinetics. In mixed culture systems, such as are used in wastewater treatment, the degree of mixing may affect the population dynamics and species selection within the reactor (Gommers et al., 1986; Yang, 1987). Liquid Mixing. It has generally been found that liquid mixing in a three-phase fluidized bed of low density particles is similar to that of beds of high density particles (Gommers et al., 1986) and is strongly dependent on the regime in which the fluidized bed is operating (Fan, 1989). Axial dispersion models are generally used to describe the degree of liquid mixing, and the presence of a gas phase has been found to have a strong effect. A near absence of axial liquid mixing (dispersed bubble regime), near complete mixing, or a mixing state somewhere between these two may occur, depending on reactor conditions (Fan, 1989). Reactor diameter has an influence on mixing, with increased influence of the gas phase on liquid mixing as the diameter increased (Gommers et al., 1986). Of interest to biofluidization is the effect of low gas flow rates, such as those found in gas-producing bioprocesses, on liquid mixing. Davison (1989) estimated the axial dispersion in both inert and fermenting threephase fluidized bed bioreactors and found that gas flow rates comparable to that produced by the fermenting organisms increased axial dispersion. Gommers et al. (1986) obtained similar results previously, though it was noted in this study that very small (1–2 mm) bubbles produced in some bioprocesses may not have as great an effect on hydrodynamics as do larger, sparged bubbles, though flow rates may be comparable. Schoutens et al. (1986a, 1986b, 1986c) found similarly small bubbles with little effect on liquid mixing in their study of a fermenting three-phase fluidized bed bioreactor. In the design of a three-phase fluidized bed bioreactor, the extent of biogas production and whether the small biogas bubbles coalesce into larger bubbles should be considered for their effect on phase mixing.

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648 Fluidization, Solids Handling, and Processing Solids Mixing and Stratification. Bioparticle size and density has been demonstrated to strongly influence solids mixing in three-phase fluidized bed bioreactors (Bly and Worden, 1990). Small, light particles were more readily drawn into bubble wakes than were large, heavy particles and are, therefore, subject to greater mixing (Tsuchiya et al., 1992). Increased gas velocity has been shown to increase axial solids dispersion; increased liquid circulation velocity, as when the static liquid to static solid bed height ratio is reduced, decreased axial solids dispersion (Tang and Fan, 1989). Fluidization regime is also important, with the larger bubbles of the coalesced and slugging regimes producing increased solids dispersion. The above is true for nonbiological fluidization as well as for biofluidization. An important difference is the changing nature of solids mixing and stratification in biofluidization as a result of biofilm formation, as was discussed in Sec. 5.3. 5.6

Mass Transfer

Mass transfer considerations are critical in any bioprocess. In typical, aerobic, suspended cell fermentations, the major concern is the oxygen transfer rate, determined by the overall mass transfer coefficient, k la, and the driving force. In three-phase biofluidization, in which the cells are immobilized as a biofilm or within carrier particles, the situation is further complicated by possible intraparticle diffusion limitations. Numerous recent studies have addressed these issues. Gas-Liquid Mass Transfer. Gas-liquid mass transfer within the three-phase fluidized bed bioreactor is dependent on the interfacial area available for mass transfer, a; the gas-liquid mass transfer coefficient, k l; and the driving force that results from the concentration difference between the bulk liquid and the bulk gas. The latter can be easily controlled by varying the inlet gas concentration. Because estimations of the interfacial area available for mass transfer depends on somewhat challenging measurements of bubble size and bubble size distribution, much of the research on increasing mass transfer rates has concentrated on increasing the overall mass transfer coefficient, k la, though several studies look at the influence of various process conditions on the individual parameters. Typical values of k la reported in the literature are listed in Table 19.

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Table 19. Typical Values of k la as Reported in the Literature

Fluid velocities can have a significant effect on k la. Increased gas velocity within the dispersed bubble regime has been found by many investigators to substantially increase k la (Tang and Fan, 1990; Karamanev et al., 1992; Nore et al., 1992; Roustan et al., 1993). Increased liquid velocity also appears to increase k la (Kang et al., 1991); Tang and Fan (1990) determined that the increase in k la was due to an increase in k l. Particle properties and solids loading have also been shown to have a significant effect on k la. Decreased particle density has been found to increase k la (Nore et al., 1992; Tang and Fan, 1990), and increased particle size also increased k la (Kang et al., 1991); the latter appears to be caused by an effect on k l , because bubble size and, therefore, interfacial area, was not affected (Sun and Furusaki, 1988). The addition of high density particles, such as those used in traditional three-phase fluidization applications, decreased k la (Kang et al., 1991; Roustan et al., 1993); increased bubble coalescence was suggested as the cause. Adding large, low density particles, of the type often used in three-phase biofluidization, was found in one study to increase k la at low solids loadings (Karamanev et al., 1992), but to decrease k la as solids holdup increased. Other investigators report a continuous decrease in k la with increased solids concentration (Sun and Furusaki, 1988; Tang and Fan, 1990); however, Roustan et al. (1993) reported little effect of low density solids holdup on k la. Contrary to studies of three-phase fluidized

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650 Fluidization, Solids Handling, and Processing bed reactors containing rigid solids, the loading of the soft solids (polyurethane) in a draft tube three-phase fluidized bed bioreactor or in a conventional three-phase fluidized bed bioreactor did not strongly affect the volumetric mass transfer coefficient (Karamanev et al., 1992). Another study suggested that the presence of low density particles (polyurethane) promoted bubble coalescence, thereby decreased interfacial area and, hence, the volumetric mass transfer coefficient (Kobayashi et al., 1990). Draft tube fluidized bed bioreactors have been very successful at increasing mass transfer rates, with increases in k la on the order of 1.5 to 3 times higher than fluidized beds without draft tubes (Karamanev et al., 1992). It has been suggested that the addition of the draft tube increases the bubble rise velocity through the increased liquid velocity within the draft tube, thereby decreasing the likelihood of bubble coalescence at the base of the reactor. The decrease in time for mass transfer appears to be offset by the larger available interfacial area for mass transfer; furthermore, many of the resulting small bubbles are entrained in the annular flow, providing additional retention time for mass transfer (Karamanev et al., 1992). The use of floating bubble breakers has been used to increase the volumetric mass transfer coefficient in a three-phase fluidized bed of glass beads (Kang et al., 1991); perhaps a similar strategy would prove effective for a bed of low density beads. Static mixers have been shown to increase k la for otherwise constant process conditions by increasing the gas holdup and, therefore, the interfacial area (Potthoff and Bohnet, 1993). Paz et al., (1993) have modified the density of calcium alginate beads by adding α-alumina to the gel in an attempt to produce beads that were more effective at breaking up bubbles in a three-phase fluidized bed bioreactor than were normal density calcium alginate beads, thereby causing an increase in interfacial area. Production of acetic acid by alginate plus alumina entrapped Acetobacter sp. cells was more than 2.5 times that of cells trapped only in alginate; the measured volumetric mass transfer coefficient was over three times greater. Understanding the effect of reactor diameter on the volumetric mass transfer coefficient is critical to successful scale up. In studies of a threephase fluidized bed bioreactor using soft polyurethane particles, Karamanev et al. (1992) found that for a classical fluidized bed bioreactor, k la could either increase or decrease with a change in reactor diameter, depending on solids holdup, but for a draft tube fluidized bed bioreactor, k la always increased with increased reactor diameter.

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Tang and Fan (1990) recommended basing volumetric mass transfer coefficient determinations on the entire axial concentration profile, thereby avoiding errors caused by end effects. Kang et al. (1991) further demonstrated the applicability of this method, successfully adopting the axial dispersion model for determination of the volumetric oxygen transfer coefficient in a three-phase fluidized bed of glass beads plus floating bubble breakers. Perhaps some of the discrepancies and contradictions seen in the various studies cited above could be explained by applying this approach to the various experimental systems. Intraparticle Mass Transfer. One way biofilm growth alters bioreactor performance is by changing the effectiveness factor, defined as the actual substrate conversion divided by the maximum possible conversion in the volume occupied by the particle without mass transfer limitation. An optimal biofilm thickness exists for a given particle, above or below which the particle effectiveness factor and reactor productivity decrease. As the particle size increases, the maximum effectiveness factor possible decreases (Andrews and Przezdziecki, 1986). If sufficient kinetic and physical data are available, the optimal biofilm thickness for optimal effectiveness can be determined through various models for a given particle size (Andrews, 1988; Ruggeri et al., 1994), and biofilm erosion can be controlled to maintain this thickness. The determination of the effectiveness factor for various sized particles with changing biofilm thickness is well-described in the literature (Fan, 1989; Andrews, 1988) Use of bioflocs rather than supported film particles will maximize the effectiveness factor for a given particle, but uneven growth of flocs can cause severe stratification in the bed. If stratification can be overcome by methods such as the use of a tapered bed to control porosity; the removal, breaking up, and recycle of biomass at the bottom of the bed; or, ideally, the use of microbial strains or species that will stop growing at a desirable floc size, such as a Zymomonas mobilis strain that stops growing at one millimeter in diameter (Scott, 1983), the use of bioflocs rather than support particles can improve reactor productivity. 5.7

Modeling

The complexity of the three-phase fluidized bed bioreactor is gradually coming under control as more sophisticated models become available. The chief need is for a model that integrates the microbial kinetics with the

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652 Fluidization, Solids Handling, and Processing reactor hydrodynamics in a dynamic fashion; both present challenging problems, and combining the two requires even greater insight. Table 20 presents the features of several models presented in the literature in recent years. It is not an exhaustive list, but demonstrates some of the advances made in modeling in recent years. These advances are expected to overcome industrial resistance to the use of these complex systems.

Table 20. Recently Developed Three-Phase Biofluidization Models I

111111111111111111111

I

I I I I I.1 I I

I IIIIIIII I I

1 1X1X euster-Botz. 1993 ang and Fan, 1999

1 = airlii 2 = inverse

1

1x1

I

I

swb I I I

airltft 3 s draft tube fluidized bed 4 - fluidized bed + stirred tank

I

A= BL CI Dt EI

ethanol production e&en1 femwntatiin phenol degradation 3,4 dichkmanaline degradntion deniwifiition

As can be seen in Table 20, only a few models have incorporated both bioreactor hydrodynamics and microbial kinetics into their scheme (Worden and Donaldson, 1987; Farag et al., 1989; Wisecarver and Fan, 1989; Petersen and Davison, 1995). Of these, only a few include experi-


653

mental validation of their simulation results, and all of these modeled the phenol degradation process (Worden and Donaldson, 1987; Wisecarver and Fan, 1989). Development of integrated models for other biofluidization applications would be beneficial to future industrial developments. Few models include the effects of in situ gas formation on the fluidization properties of the reactors; this improvement, along with improvements in other areas, such as inclusion of improved structured models of microbial kinetics or inclusion of maintenance energy requirements or the effects of suspended cells on the reaction rate, might produce more accurate models, though it is unclear at this point whether the increased complexity would be justified. 5.8

Scale Up

As mentioned previously, the design of a three-phase fluidized bed bioreactor is complex, even more so than the design of a nonbiological three-phase fluidized bed reactor. Scale-up to commercial size is relatively common for wastewater treatment, but because of the increased constraints in biosynthetic processes, fewer examples exist for these. Several are listed in Table 21. As process development techniques continue to improve as a result of advances in the understanding of engineering fundamentals, computing power, and analytical methods, the daunting task of designing a commercial scale three-phase fluidized bed bioreactor becomes less intimidating and is expected to be undertaken more frequently in the future. The traditional, linear scale-up approach, proceeding from kinetic studies at lab scale, determining pilot plant specific correlations in small pilot plants, building a large demonstration unit based on pilot plant experiments then developing a model specific to the large demonstration unit, and, finally, building the commercial reactor and only then tuning the correlations for the final plant, is being replaced by a more interactive approach, in which a unified model that considers kinetics, thermodynamics, hydrodynamics, and hardware is reconciled to data at all scales. Not only does this allow more confidence in the final design, it also results in a better understanding of the process (Tarmy and Coulaloglou, 1992). An example of the application of a unified model to the design of a three-phase, fluidized bed bioreactor is the scale down, scale up procedure. A model of the full scale reactor is developed, then is used to design

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654 Fluidization, Solids Handling, and Processing a scaled-down laboratory reactor. The laboratory reactor is built and used for a series of experiments, the data from which are used to validate the model. If the model is validated, it can then be used as the basis for a full scale design. Schoutens et al. (1986a, 1986b, 1986c) used this approach to show that a fluidized bed reactor was the preferred design for the production of butanol and isopropanol by Clostridium spp. immobilized in calcium alginate beads (Schoutens et al., 1986a). The model, developed for a 50-65 m3 reactor, was shown to be accurate in predicting the results in 10-15 dm3 reactors, thereby confirming its reliability for large scale design, according to the authors (Schoutens et al., 1986c). Table 21. Examples of Pilot and Large Scale Applications of Three-Phase Biofluidization to Fermentation Processes

Another, more traditional, approach to scale-up is application of similarity through dimensional analysis. One comparison of several such design criteria (biological, dynamic, and geometric) for biological fluidized bed reactors found that of the parameters tested, the sludge retention time (STR), the Peclet number, and the particle Reynolds number were most suitable for scale-up (Ozturk et al., 1994). Because of constraints on particle size and material, and fluid viscosity, it is difficult to maintain similarity in all relevant dynamic similarity criteria simultaneously. Conservative measures, such as using low estimates of such values as the overall mass transfer coefficient (Ryhiner et al., 1988), are suggested, in any case.

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Time constants for the process can provide relevant information for scale up. A comparison of the conversion time constant to the time constant for mixing allows determination of whether mixing will affect the process (Gommers et al., 1986; Schoutens et al., 1986a, 1986b, 1986c). If mixing is incomplete or the time constant for mixing is much larger than the time constant for conversion, conditions will differ at different heights in the reactor and biological stratification, which may or may not be desirable, may occur (Gommers et al., 1986). Few other design examples exist in the literature. Schneeberg (1994) provides one in which a fluidized bed upgrade is added to an existing paper mill wastewater treatment plant. Badot et al. (1994) also gave a brief example of design for both a pilot scale and industrial scale reactor for various wastewater treatment applications. Andrews (1988) provided another example of aerobic wastewater treatment focusing on bioparticle optimization. 5.9

Process Strategy

Before the details of a particular reactor are specified, the biochemical engineer must develop a process strategy that suits the biokinetic requirements of the particular organisms in use and that integrates the bioreactor into the entire process. Reactor costs, raw material costs, downstream processing requirements, and the need for auxiliary equipment will all influence the final process design. A complete discussion of this topic is beyond the scope of this chapter, but a few comments on reactor choice for particular bioprocesses is appropriate. For those bioprocesses that follow Monod kinetics without substrate inhibition, a reactor configuration approaching plug flow behavior will give the highest rates of conversion. In three-phase fluidization, this means a high height to diameter ratio, within power consumption limits. The use of multiple stages is one possibility for overcoming the power consumption limitation on reactor height and would have additional benefits for bioprocesses in which separation of biomass growth and product generation is desirable (Fan, 1989). If substrate inhibition exists, a well-mixed bioreactor is desirable. Mixing in three-phase fluidized bed bioreactors can be increased by adding an external recycle loop, by inserting a draft tube in the reactor, or by decreasing the height to diameter ratio.

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656 Fluidization, Solids Handling, and Processing Conversely, product inhibition suggests the use of a reactor that emulates plug flow, as for the Monod kinetics without substrate inhibition discussed above. Alternatively, in situ product removal may alleviate product inhibition and improve productivity (Yabannavar and Wang, 1991). A bioparticle reactor for simultaneous fermentation/adsorption of fermentation or enzyme reaction products has been suggested and demonstrated by several workers (van der Wielen et al., 1990; Davison and Scott, 1992; Kaufman et al., 1995). Such a reactor makes use of the particle stratification observed in fluidized bed reactors, according to particle size and density. This allows a denser adsorbent particle to be added at the top of the reactor, from which it falls through the fluidized bed of biocatalysts, absorbing product as it passes. The dense beads are removed from the base of the bed and may be regenerated by product recovery for further use. By removing the product, which may inhibit the biological reactions taking place in the reactor, productivity is improved. When mixed cultures are involved, the reaction kinetics may be complicated by interspecies relationships. For example, in the treatment of various waste waters, a Monod-type kinetics relating removal rate to organic loading rate has been verified (Converti et al., 1993). As expected, when the loading rate was increased or the residence time was decreased, either of which increase the substrate concentration, removal rate increased (Converti et al., 1993); however, this increased removal rate came at the cost of decreased removal efficiency. The initial biodegradative reaction (acidogenic) increased in rate faster than did the final reaction (methanogenic), resulting in a deleterious accumulation of organic acids. This had to be monitored and ameliorated through the addition of base to the reactor to protect the methanogenic stage (Converti et al., 1993). For situations, such as the example above, where mixed cultures of degradative organisms have different optimal growth conditions, multiple stages would be desirable (Converti et al., 1993). This approach has been successfully demonstrated and, with the addition of a recycle stream from the second stage to the first, has been shown to have the additional costsaving benefit of reducing operating costs by decreasing the amount of caustic required for pH control, the major operational cost for full-scale anaerobic treatment plants (Romli et al., 1994). Another approach to multiple environmental optima for various species is the stratified bed suggested by Yang (1987), shown in Fig. 12. Use of support beads of various terminal velocities and a tapered bed can result in little or no axial

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solid mixing; organisms would then be allowed to colonize the particular reactor region in which environmental conditions were optimal for each species involved. G

G

L: Liquid G: Gas

Stream Stream

P: Prodlt:t

L

.Immobilized o Gas

Cell Partcles

Bubbles

Stream

Figure 12. Tapered-bed three-phase fluidized bed bioreactor.

5.10 Novel Reactors Magnetically stabilized fluidized beds, in which a magnetic field is applied to the fluidized bed of magnetic particles, stabilizing the bed, have been suggested for several different bioprocesses (Hu and Wu, 1987; Siegell, 1987; Bramble et al., 1990; Terranova and Burns, 1991; Weng et al., 1992). Among the advantages cited for such a bioreactor include low pressure drops and ease of solids handling as expected in fluidization, yet somewhat ordered bed structure with little or no solids mixing, depending on flow regime, as expected in a packed bed, so that particle residence time can be tightly controlled. The stabilized bed allows higher flow rates without particle elutriation than could be used in conventional fluidization. Possible fluidization regimes described in the literature include particulate fluidization under weak magnetic fields, chain fluidization under increased magnetic fields, and, finally, magnetically condensed, in

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658 Fluidization, Solids Handling, and Processing which the bed is no longer fluidized. For the particulate and chain regimes, bed height has been shown to decrease as the field strength is increased. Gas holdup was affected by magnetic field strength, as well, increasing with increasing field strength (Hu and Wu, 1987), probably because magnetically stabilized particles had a greater tendency to break bubbles apart (Weng et al., 1992), and smaller bubbles have a slower rise velocity, resulting in greater gas holdup. Gas-liquid mass transfer was increased for the same reason (Weng et al., 1992). Radial distribution of gas and liquid flows tended to flatten with increased magnetic field (Hu and Wu, 1987; Weng et al., 1992). Several applications of magnetically stabilized fluidized beds for bioprocessing have been demonstrated. Continuous cell suspension processing has been demonstrated in a magnetically stabilized fluidized bed that acted like a moving depth filter from which the solid matrix could be continuously removed to avoid clogging (Terranova and Burns, 1991); though this was demonstrated in a two-phase system, similar three-phase applications could be envisioned. Continuous ethanol fermentation has been performed in a three-phase magnetic fluidized bed (Weng et al., 1992). Production of caffeine and theobromine by plant cells in alginate magnetized with magnetite (Fe 3O4) has also been demonstrated (Bramble et al., 1990). Other applications are likely to be demonstrated in the future. Several improvements in the conventional airlift reactor have been suggested. Inverse fluidized bed bioreactors, based on airlift bioreactor principles and shown in Fig. 13, have been developed by several workers (Garnier et al., 1990; Nikolov and Karamanev, 1987, 1990). Low density, biofilm-covered particles are floated in the annular region of the airlift reactor; as the biofilm thickness increases, the particle density also increases and the particle bed expands downward. When the density is sufficiently high so that the particles reach the bottom of the bed, the particles are swept into the airlift sparged section in the reactor center where hydrodynamic forces strip some of the biofilm from the particles which then travel to the top of the reactor and return to the annular region. This cyclic mixing pattern has been shown to maintain a uniform biofilm thickness (Garnier et al., 1990). By adding a contained fluidized bed of coarser, denser particles in the airlift section, Nikolov and Karamanev (1987), improved attrition of biofilm from the particles. This reactor was superior to a conventional airlift reactor that used suspended cells for aerobic wastewater treatment by a mixed bacterial culture and for ferrous iron oxidation by Thiobacillus ferroxidans.

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L

.0.

659

G

~.o.

~~.Oo~~..ci' .=~.

00...00. .0

o .

.

..0

00 0.00 0 .-Oeo . ..0 o 0 0 ~o . 0 ..000. 0 0 0

G

L: Liquid G : Gas

Stream. Stream

P: ProdlJ:t

Immobilized o Gas

Cell PartK:les

Bubbles

Stream

Figure 13. Inverse three-phase fluidized bed bioreactor.

In order to increase radial mixing in airlift reactors, Merchuk et al. (1993) placed helical flow promoters (helical baffles) into the upper annular region, causing secondary flows that not only increased radial mixing, thereby improving light (for plant cell culture), heat, and mass transfer, but also enhanced the fluidization of solid particles in the reactor by generating swirling flow at the base of the reactor. The dually injected turbulent separation (DITS) three-phase fluidized bed bioreactor described by van der Jagt et al. (1993) separated large and small soil particles for optimal simultaneous bioremediation. Air and recycled liquid were simultaneously injected into the reactor base, fluidizing coarse, heavy soil particles in the lower region of the reactor and suspending fine particles in the upper region. Because the small particles in a contaminated soil carry a disproportionate amount of the pollutant, they require greater time for efficient biodegradation of the contaminant. By exploiting the stratification caused by fluidization, this reactor design allows large particles to be drawn off separately from the small particles, which are recycled to the bottom of the reactor and thus receive the longer treatment they require. Reactor size and cost is kept at a minimum, because each portion is treated only as long as absolutely required.

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660 Fluidization, Solids Handling, and Processing A continuous centrifugal bioreactor, in which cells are fluidized in balance with centrifugal forces, has been designed to allow high density cell cultivation and superior aeration without elutriation of the suspended cells (van Wie et al., 1991). Reactor performance was hampered by elutriation of biomass by evolved gas in an anaerobic fermentation, indicating that it may not be suitable in its present state for three-phase fermentations. Immobilization of the cells on denser particles may overcome this problem. In order to simultaneously achieve the advantages of plug flow operation (controlled residence time, reduced product inhibition) along with the advantages of intense mixing in each consecutive stage (good transport properties), the Blenke-Cascade reactor has been designed at the University of Hohenheim (Kottke et al., 1991). This reactor is a heavily baffled tower divided into many chambers by the baffles. Each chamber is well-mixed by fluidizing gas passing up the tower; the solid particles (biocatalysts) and liquid may flow co- or countercurrently. Several researchers have suggested contained fluidized bed for bioprocesses; such a reactor is depicted in Fig. 14. Kalogerakis and Behie (1995) have designed a three-phase/two region bioreactor for vaccine production. The bioparticles (microcarriers for animal cell culture) are contained in one region of the reactor and are kept in suspension by gentle mechanical agitation. An overall liquid flow is imposed on the whole reactor by the action of sparging in a separate central region, resulting in something of a hybrid of an airlift reactor and agitated reactor. Because the sparged section provides excellent aeration, the outer cell-containing region requires only mild agitation. Naouri et al. (1991) described another contained fluidized bed, the so-called high compacting multiphasic reactor (HCMR), which they used for malic and lactic acid fermentations for wine improvement. Bioparticles were contained within a tapered region and liquid was rapidly recycled through this region by pumping; improved liquid/solid contact was cited as the advantage of this reactor. Gas logging, the adherence of small bubbles to particles, causing them to rise to the surface in the reactor and form an inefficient packed bed with poor mass transfer properties, can be a problem in various fermentations and in wastewater treatment. A double entry fluidized bed reactor has been developed with simultaneous top (inverse) and bottom (conventional) inlets to overcome this problem (Gilson and Thomas, 1993).

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Effluent is drawn off from a side port. Under conditions in which a conventional three-phase fluidized bed bioreactor operated in an unstable manner because of gas logging, the new bioreactor converted glucose to ethanol at a 27% higher rate. Saccharomyces cerevisiae was grown in alginate beads for this reaction.

Retaining Grid

L

L: Liquid G : Gas

Stream Stream

P; Prod~t

.I o

mmobilized Gas

Cell

PartK:les

Bubbles

Stream

Figure 14. Contained three-phase fluidized bed bioreactor.

5.11 Economics The economics of ethanol fermentations have been examined by several researchers (Busche and Allen, 1989; Busche et al., 1992; Harshbarger et al., 1995). Though the economic feasibility of the ethanol fermentation is dependent on the price of competitive petrochemical routes (and is currently unfavorable), the application of three-phase fluidized bed bioreactor technology improves the economics over those for conventional reactors by reducing operating costs through improved conversion and yield, by lowering utilities costs, and by reducing capital costs. Similar improvements were found when three-phase fluidized bed bioreactor technology was applied to the acetone-butanol-ethanol fermentation, with the addition of a membrane process to remove product and

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662 Fluidization, Solids Handling, and Processing prevent product inhibition (Qureshi and Maddox, 1992). Another ABE fermentation scheme that included product removal by the inclusion of a fourth, extractive phase, showed outstanding improvements in process economics, reducing projected product costs by more than half (Busche and Allen, 1989); this improvement was not sufficient to beat the price for petrochemically derived butanol, but is an indication that as petroleum prices rise, three-phase biofluidization has a definite place in bioprocess design. These improvements in process economics are likely to be applicable to other fermentation processes. For waste treatment rather than fermentation for product formation, again few examples of process economics exist in the literature. Those that do, favor fluidization. Badot et al. (1994) described an industrial prototype fluidized bed reactor that competed favorably on an economical basis with activated sludge processes for treating carbon pollution and was estimated to be economically comparable to fixed bed processes for denitrification. Schneeberg (1994) described the successful and economically-sound implementation of fluidization as an upgrade to an existing wastewater treatment plant. The restricted space available for extension of the wastewater plant made fluidization particularly advantageous in this case. 5.12 Summary The complexity of the interaction of reactor hydrodynamics and performance with the biological metabolic processes is the major deterrent to widespread use of three-phase fluidization in biological processes. Most commercial applications of three-phase fluidization have thus far been in the area of wastewater treatment. Commercial fermentation and animal cell culture applications exist, but much of the work in this area remains at laboratory scale. In fermentation or cell culture systems, the solid phase is typically much lower in density than that found in traditional three-phase fluidized systems, and reactor hydrodynamics and transport properties can vary dramatically from those well-studied systems. As a greater understanding of reactor fundamentals for the new applications is obtained, more successful commercial applications of three-phase biofluidization for fermentations and cell culture are likely. Fortunately, the last few years have seen great strides in understanding the complex relationships of biofluidization, allowing increased confidence in design of large scale systems based on laboratory data.

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ACKNOWLEDGMENT The authors would like to express their appreciation to researchers at BHP in Australia for providing information on the recent developments in smelting reduction of iron ore.

NOTATIONS a B

= =

CD do dp e Fr Fz g ∆Hr ka kl n r SH T ui

= = = = = = = = = = = = = = =

ul ut u ts x z

= = = = =

Gas-liquid interfacial area Ratio of buoyant density of the biofilm to the buoyant density of the support Drag coefficient Diameter of biofilm-free particle Diameter of a particle Exponent on Reynolds number coefficient correlation Radiation heat flux in the r direction Radiation heat flux in the z direction Gravitational acceleration Heat of reaction Absorption coefficient Gas-liquid mass transfer coefficient Richardson-Zaki index Radial coordinate Source term for enthalpy equation Temperature Extrapolated liquid velocity as ε→1 for liquid-solid fluidized bed Superficial liquid velocity Terminal settling velocity of a single particle Terminal settling velocity of a biofilm-free particle Ratio of film volume to support volume Axial coordinate

Greek Symbols ε ρ bw

= =

Void fraction Wet density of biofilm

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664 Fluidization, Solids Handling, and Processing ρf ρp ρs

= = =

σ

=

Density of fluid Apparent density of bioparticle Density of nonporous particle or effective density of nearly nonporous particle Stefan-Boltzmann constant

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Macy, J. M., Lawson, S., and DeMoll-Decker, H., “Bioremediation of Selenium Oxyanions in San Joaquin Drainage Water Using Thauera selenatis in a Biological Reactor System,” Appl. Microbiol. Biotechnol., 40:588 (1993) Mäkinen, P. M., Theno, T. J., Ferguson, J. F., Ongerth, J. E., and Puhakka, J. A., “Chlorophenol Toxicity Removal and Monitoring in Aerobic Treatment: Recovery from Process Upsets,” Environ. Sci. Technol. 27:1434 (1993) Matsui, S., Yamamoto, R. I., Tsuchiya, Y., and Inanc, B., “The Kinetics of Glucose Decomposition with Sulfate Reduction in the Anaerobic Fluidized Bed Reactor,” Wat. Sci. Tech., 28(2):135 (1993) McDonough, T. J., “Recent Advances in Bleached Chemical Pulp Manufacturing Technology, Part 1: Extended Delignification, Oxygen Delignification, Enzyme Applications, and ECF and TCF Bleaching,” Tappi J., 78(3):55 (1995) McKinney, R. W. J., “Wastepaper Preparation and Contamination Removal,” Technol. of Paper Recycling, (R. W. J. McKinney, ed.), p. 48, Blackie Academic & Professional, London (1995) Meadows, D. G., “The Pulp Mill of the Future: 2005 and Beyond,” Tappi J., 78(10):55 (1995) Merchuk, J. C., Ladwa, N., and Bulmer, M., “Improving the Airlift Reactor: The Helical Flow Promoter,” 3rd Int. Conf. on Bioreactor and Bioprocess Fluid Dynamics, (A. W. Nienow, ed.), p. 61, Information Press, Ltd., Oxford (1993) Merchuk, J. C., Ben-Zvi (Yona), S., and Niranjan, K., “Why Use Bubble-Column Bioreactors?,” Trends Biotechnol., 12:501 (1994) Mol, N., Kut, O. M., and Dunn, I. J., “Adsorption of Toxic Shocks on Carriers in Anaerobic Biofilm Fluidized Bed Reactors,” Wat. Sci. Tech., 28(2):55 (1993) Mollah, A. H., and Stuckey, D. C., “Maximizing the Production of AcetoneButanol in an Alginate Bead Fluidized Bed Reactor Using Clostridium acetobutylicum,” J. Chem. Tech. Biotechnol., 56:83 (1993) Muroyama, K., and Fan, L. -S., “Fundamentals of Gas-Liquid-Solid Fluidization,” AIChE J., 31:1 (1985) Myška, J., and Švec, J., “The Distributive Properties of a Fluidized Bed with Biomass,” Wat. Res., 28:1653 (1994) Nagamune, T., Endo, I., Kato, N., Nishimura, M., and Kobayashi, T., “The Effect of Cultivation Conditions on the Penicillin Production Using a Urethane Foam-Supported Penicillium chrysogenum,” Bioproc. Eng., 3:173 (1988) Naouri, P., Bernet, N., Chagnaud, P., Arnaud, A., and Galzy, P., “Bioconversion of L-Malic Acid into L-Lactic Acid Using a High Compacting Multiphasic Reactor (HCMR),” J. Chem. Tech. Biotechnol., 51:81 (1991)

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674 Fluidization, Solids Handling, and Processing Nelson, T. B., and Skaates, J. M., “Attrition in a Liquid Fluidized Bed Bioreactor,” Ind. Eng. Chem. Res., 27:1502 (1988) Nguyen, V. T., and Shieh, W. K., “Continuous Ethanol Fermentation Using Immobilized Yeast in a Fluidized Bed Reactor,” J. Chem. Tech. Biotechnol., 55:339 (1992) Nikolov, L. N., and Karamanev, D., “Experimental Study of the Inverse Fluidized Bed Biofilm Reactor,” Can. J. Chem. Eng., 65:214 (1987) Nikolov, L., and Karamanev, D. G., “The Inverse Fluidized Bed Biofilm Reactor: A New Laboratory Scale Apparatus for Biofilm Research,” J. Ferm. Bioeng. 69:265 (1990) Nore, O., Briens, C., Margaritis, A., and Wild, G., “Hydrodynamics, Gas-Liquid Mass Transfer and Particle-Liquid Heat and Mass Transfer in a ThreePhase Fluidized Bed for Biochemical Process Applications,” Chem. Eng. Sci., 47:3573 (1992) Nutt, W. E., Griggs, B. F., Eachus, S. W., and Pikulin, M. A., “Developing an Ozone Bleaching Process,” Tappi J., 76(3):115 (1993) Oeters, F., “Mixing with Heat and Mass Transfer in Metallurgical Systems,” Steel Res., 60:185 (1989) Oeters, F., Zhang, L., and Steffen, R., “Considerations on Process Technologies of Smelting Reduction Including Scrap Melting,” Turkdogan Symp. Proc., p. 23 (1994) Oeters, F., and Xie, H., “A Contribution to the Theoretical Description of MetalSlag Reaction Kinetics,” Steel Res., 66:409 (1995) Ogata, M., Marumoto, Y., Oh-I, K., Shimizu, S., and Katoh, S., “Continuous Culture of CHO-K1 Cells Producing Thrombomodulin and Estimation of Culture Conditions,” J. Ferm. Bioeng., 77:46 (1994) Ogawa, Y., Katayama, H., Hirata, H., Tokumitsu, N., and Yamauchi, M., “Slag Foaming in Smelting Reduction and Its Control with Carbonaceous Materials,” ISIJ Int., 32:87 (1992) Ogawa, Y., Huin, D., Gaye, H., and Tokumitsu, N., “Physical Model of Slag Foaming,” ISIJ Int., 33:224 (1993) Oh, D. K., Hyun, C. K., Kim, J. H., and Park, Y. H., “Production of Penicillin in a Fluidized-Bed Bioreactor: Control of Cell Growth and Penicillin Production by Phosphate Limitation,” Biotechnol. Bioeng., 32:569 (1988) Ozturk, I., Turan, M., and Idris, A. H., “Scale-Up and Biomass Hold-Up Characteristics of Biological Fluidized Bed Reactors,” Wat. Sci. Tech., 29(10–11):353 (1994) Paulonis, M. A., and Krishnagopalan, A., “Kappa Number and Overall Yield Calculation Based on Digester Liquor Analysis,” Tappi J., 71(11):185 (1988)

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Paz, E. D. D., Santana, M. H. A., and Eguchi, S. Y., “Enhancement of the Oxygen Transfer in a Circulating Three-Phase Fluidized Bed Bioreactor,” Appl. Biochem. Biotechnol., 39/40:455 (1993) Petersen, J. N., and Davison, B. H., “Development of a Predictive Description of an Immobilized-Cell, Three-Phase, Fluidized-Bed Bioreactor,” Biotechnol. Bioeng., 46:139 (1995) Petrov…i…, J., Hvala, N., Bitenc, A., and Strm…nik, S., “Low-Cost Steam Consumption Control System for Batch Pulp Cooking,” Control Eng. Practice, 3(3):357 (1995) Petrozzi, S., Kut, O. M., and Dunn, I. J., “Protection of Biofilms Against Toxic Shocks by the Adsorption and Desorption Capacity of Carriers in Anaerobic Fluidized Bed Reactors,” Bioproc. Eng., 9:47 (1993) Petrozzi, S., and Dunn, I. J., “Biological Cyanide Degradation in Aerobic Fluidized Bed Reactors: Treatment of Almond Seed Wastewater,” Bioproc. Eng., 11:29 (1994) Pfeffer, J. T., and Suidan, M. T., “Continuous Processing of Toxic Organics in a Fluidized-Bed GAC Reactor Employing Carbon Replacement,” Biotechnol. Bioeng., 33:139 (1989) Potthoff, M., and Bohnet, M., “Influence of Solids Concentration and Static Mixers on Fluid Dynamics in Three-Phase Fluidized Bed Bioreactors,” Chem. Eng. Technol., 16:147 (1993) Probstein, R. F., and Hicks, R. E., Synthetic Fuels, McGraw-Hill, New York, (1982) Puhakka, J. A., Herwig, R. P., Koro, P. M., Wolfe, G. V., and Ferguson, J. F., “Biodegradation of Chlorophenols by Mixed and Pure Cultures from a Fluidized-Bed Reactor,” Appl., Microbiol. Biotechnol., 42:951 (1995) Qureshi, N., and Maddox, I. S., “Reactor Design for the ABE Fermentation Using Cells of Clostridium acetobutylicum Immobilized by Adsorption onto Bonechar,” Bioproc. Eng., 3:69 (1988) Qureshi, N., and Maddox, I. S., “A Mathematical Model of a Fluidized Bed Reactor for the Continuous Production of Solvents by Immobilized Clostridium acetobutylicum,” J. Chem. Tech. Biotechnol., 48:369 (1990) Qureshi, N., and Maddox, I. S., “Application of Novel Technology to the ABE Fermentation Process: An Economic Analysis,” Appl. Biochem. Biotechnol., 34/35:441 (1992) Racher, A. J., and Griffiths, J. B., “Investigation of Parameters Affecting a Fixed Bed Bioreactor Process for Recombinant Cell Lines,” Cytotechnol., 13:125 (1993) Ramsay, B. A., Wang, D., Chavarie, C., Rouleau, D., and Ramsay, J. A., “Penicillin Production in an Inverse Fluidized Bed Bioreactor,” J. Ferm. Bioeng., 72:495 (1991)

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676 Fluidization, Solids Handling, and Processing Reese, J., Jiang, P., and Fan, L. S., “Bubble Characteristics in Three-Phase Systems used for Pulp and Paper Processing,” Chem. Eng. Sci., 51:2501 (1996) Reiter, M., Buchacher, A., Blüml, G., Zach, N., Steinfellner, W., Schmatz, C., Gaida, T., Assadian, A., and Katinger, H., “Production of the HIV-1 Neutralising Human Monoclonal Antibody 2F5: Stirred Tank Versus Fluidized Bed Culture,” Animal Cell Technology, (R. E. Spier, J. B. Griffith, and W. Berthold, eds.), p. 333, Butterworth-Heinemann, Oxford (1994) Richardson, J. F., and Zaki, W. N., “Sedimentation and Fluidization Part I,” Trans. Inst. Chem. Engrs., 32:35 (1954) Rickard, J. C., “The 2020 Mill-Utility Requirements for Market Kraft Pulp,” AIChE Symp. Series: Advances in Forest Products Environmental and Process Eng., The 1993 Forest Products Symp., (C. L. Verrill, ed.), 90(302):167, AIChE, New York (1994) Rihar, M., “The Software Simulator as an Effective Tool for Testing Control Algorithms,” Simulation, 63(1):6 (1994) Romenets, V. A., “Process of Liquid Phase Reduction of Iron: Development and Implementation,” Steel in the USSR, 20:366 (1990) Romli, M., Greenfield, P. F., and Lee, P. L., “Effect of Recycle on a Two-Phase High-Rate Anaerobic Wastewater Treatment System,” Wat. Res., 28:475 (1994) Roustan, M., Capdeville, B., and Bastoul, D., “Influence of the Hydrodynamic Behaviour of a Three-Phase Fluidised Bed Used for Effluent Treatment,” 3rd Int. Conf. on Bioreactor and Bioprocess Fluid Dynamics, (A. W. Nienow, ed.), p. 323, Information Press, Ltd., Oxford (1993) Ruggeri, B., Caire, G., Specchia, V., Sassi, G., Bosco, F., and Gianetto, A., “Determination of Optimal Biofilm Activity in a Biological Fluidized Bed (BFB) Reactor,” Wat. Sci. Tech., 29(10-11):347 (1994) Ryhiner, G., Petrozzi, S., and Dunn, I. J., “Operation of a Three-Phase Biofilm Fluidized Sand Bed Reactor for Aerobic Wastewater Treatment,” Biotechnol. Bioeng., 32:677 (1988) Sajc, L., Obradovic, B., Vukovic, D., Bugarski, B., Grubisic, D., and VunjakNovakovic, G., “Hydrodynamics and Mass Transfer in a Four-Phase External Loop Air Lift Bioreactor,” Biotechnol. Prog., 11:420 (1995) Samejima, H., Nagashima, M., Azuma, M., Noguchi, S., and Inuzuka, K., “Semicommercial Production of Ethanol using Immobilized Microbial Cells,” Annals New York Academy of Sciences, 434:394 (1984) Sanz, I., and Fdz-Polanco, F., “Low Temperature Treatment of Municipal Sewage in Anaerobic Fluidized Bed Reactors,” Wat. Res., 24:463 (1990)

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678 Fluidization, Solids Handling, and Processing Shirai, Y., Yamaguchi, M., Kobayashi, A., Nishi, A., Nakamura, H., and Murakami, H., “Change in Growth Kinetics of Hybridoma Cells Entrapped in Collagen Gel Affected by Alkaline Supply,” Cytotechnol., 14:129 (1994) Shu, C. H., and Yang, S. T., “Effect of Particle Loading on GM-CSF Production by Saccharomyces cerevisiae in a Three-Phase Fluidized Bed Bioreactor,” Biotechnol. Bioeng., 51:229 (1996) Siegel, M. H., and Robinson, C. W., “Applications of Airlift Gas-Liquid-Solid Reactors in Biotechnology,” Chem. Eng. Sci., 47:3215 (1992) Siegell, J. H., “Liquid-Fluidized Magnetically Stabilized Beds,” Powder Tech., 52:139 (1987) Silbiger, E., and Freeman, A., “Continuous ∆1 -Hydrocortisone Dehydrogenation with in Situ Product Recovery,” Enzyme Microb. Technol., 13:869 (1991) Smook, G. A., Handbook for Pulp & Paper Technologists, 2nd ed., Angus Wilde Publications, Vancouver (1992) Sreekrishnan, T. R., Ramachandran, K. B., and Ghosh, P., “Effect of Operating Variables on Biofilm Formation and Performance of an Anaerobic FluidizedBed Bioreactor,” Biotechnol. Bioeng., 37:557 (1991) Steffen, R., “Direct Reduction and Smelting Reduction - An Overview,” Steel Res., 60:96 (1989) Stenuf, T. J., and Unbehend, J. E., “Hydrodynamics of Fiber Suspensions,” Encyclopedia of Fluid Mechanics, Slurry Flow Technol., (N. Cheremisinoff, ed.), 5:291, Gulf Publishing Co., Houston (1986) Suh, I. S., Schumpe, A., Deckwer, W. -D., “Xanthan Production in Bubble Column and Air-Lift Reactors,” Biotechnol. Bioeng., 39:85 (1992) Sun, Y., and Furusaki, S., “Mean Bubble Diameter and Oxygen Transfer Coefficient in a Three-Phase Fluidized Bed Bioreactor,” J. Chem. Eng. Jpn., 21:20 (1988) Sun, Y., and Furusaki, S., “Continuous Production of Acetic Acid Using Immobilized Acetobacter aceti in a Three-Phase Fluidized Bed Bioreactor,” J. Ferm. Bioeng., 69:102 (1990) Takahashi, K., Muroya, M., Kondo, K., Hasegawa, T., Kikuchi, I., and Kawakami, M., “Post Combustion Behavior in In-Bath Type Smelting Reduction Furnace,” ISIJ Int., 32:102 (1992) Takashiba, N., Nira, M., Kojima, S., Take, H., and Yoshikawa, F., “Development of the Post Combustion Technique in Combined Blowing Converter,” Tetsu to Hagane, 75:89 (1989) Tanabe, H., Takahashi, K., Ariyama, Y., Kitagawa, T., Ozeki, A., and Yamaga, M., “The Primary Stage of the Development of Smelting Reduction Process in NKK,” AIME Ironmaking Conf. Proc., p. 89 (1989)

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Tanaka, H., Enoki, N., Kaneki, N., Sakai, H., Shimada, K., and Hara, H., “A Three-Phase Fluidized Bed Fuel Cell,” J. Electrochem. Soc., 137:2798 (1990) Tanemura, K., Kida, K., Ikbal, Matsumoto, J., and Sonoda, Y., “Anaerobic Treatment of Wastewater with High Salt Content from a Pickled-Plum Manufacturing Process,” J. Ferm. Bioeng. 77:188 (1994a) Tanemura, K., Kida, K., Teshima, M., and Sonoda, Y., “Anaerobic Treatment of Wastewater from a Food-Manufacturing Plant with a Low Concentration of Organic Matter and Regeneration of Usable Pure Water,” J. Ferm. Bioeng., 77:307 (1994b) Tang, W. T., and Fan, L. -S., “Steady State Phenol Degradation in a Draft-Tube, Gas-Liquid-Solid Fluidized-Bed Bioreactor,” AIChE J., 33:239 (1987) Tang, W. T., Wisecarver, K., and Fan, L. S., “Dynamics of a Draft Tube GasLiquid-Solid Fluidized Bed Bioreactor for Phenol Degradation,” Chem. Eng. Sci., 42:2123 (1987) Tang, W. T., and Fan, L. -S., “Hydrodynamics of a Three-Phase Fluidized Bed Containing Low-Density Particles,” AIChE J., 35:355 (1989) Tang, W. -T., and Fan, L. -S., “Gas-Liquid Mass Transfer in a Three-Phase Fluidized Bed Containing Low Density Particles,” Ind. Eng. Chem. Res., 29:128 (1990) Tarmy, B. L., Chang, M., Coulaloglou, C. A., and Ponzi, P. R., “The Three-Phase Hydrodynamic Characteristics of the EDS Liquefaction Reactors: Their Development and use in Reactor Scaleup,” Proc. 8th Int. Symp. Chem. Reaction Eng., 30:239 (1984) Tarmy, B. L., and Coulaloglou, C. A., “Alpha-Omega and Beyond - Industrial View of Gas/Liquid/Solid Reactor Development,” Chem. Eng. Sci., 47:3231 (1992) Terashima, S., Kamihira, M., Ogawa, T., Ohno, M., Iijima, S., and Kobayashi, T., “Continuous Production of Human Erythropoietin by Immobilized Recombinant L-929 Cells,” J. Ferm. Bioeng., 77:52 (1994) Terranova, B. E., and Burns, M. A., “Continuous Cell Suspension Processing Using Magnetically Stabilized Fluidized Beds,” Biotechnol. Bioeng., 37:110 (1991) Tikka, P. O., Virkola, N. E., Pursiainen, S. A., and Haemaelae, I. T., “Process Chemistry and Control of Rapid-Displacement Heating,” Tappi J., 71(2):51 (1988) Tong, C. C., and Fan, L. S., “Concentration Multiplicity in a Draft Tube FluidizedBed Bioreactor Involving Two Limiting Substrates,” Biotechnol. Bioeng., 31:24 (1988)

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680 Fluidization, Solids Handling, and Processing Toseland, B. A., Brown, D. M., Zou, B. S., and Dudukovi…, M., “Flow Patterns in a Slurry-Bubble-Column Reactor Conditions,” Trans. Inst. Chem. Engrs., 73:297 (1995) Trinet, F., Heim, R., Amar, D., Chang, H. T., and Rittmann, B. E., “Study of Biofilm and Fluidization of Bioparticles in a Three-Phase Liquid-FluidizedBed Reactor,” Wat. Sci. Tech., 23:1347 (1991) Tseng, S. -K., and Lin, M. -R., “Treatment of Organic Wastewater by Anaerobic Biological Fluidized Bed Reactor,” Wat. Sci. Tech., 29(12):157 (1994) Tsuchiya, K., Song, G. H., Tang, W. T., and Fan, L. S., “Particle Drift Induced by a Bubble in a Liquid-Solid Fluidized Bed with Low Density Particles,” AIChE J., 38:1847 (1992) Tsujino, R., Mukai, T., Hirai, M., Nakamura, K., and Harada, T., “Investigation of the Effect on Post Combustion in LD Converter (Study of Post Combustion in LD Converter-III),” Trans. ISIJ, 25:B-294 (1985) Turvey, R. W., “Chemical use in Recycling,” Technol. of Paper Recycling, (R. W. J. McKinney, ed.), p. 130, Blackie Academic & Professional, London (1995) Tzeng, J. W., Study of Fluidized Bed Reactors - Fluid Dynamics and Bioreactor Appl., Doctoral Dissertation, The Ohio State University (1991) Tzeng, J. W., Fan, L. S., Gan, Y. R., and Hu, T. T., “Ethanol Fermentation Using Immobilized Cells in a Multistage Fluidized Bed Bioreactor,” Biotechnol. Bioeng., 38:1253 (1991) Utigard, T. A., and Zamalloa, M., “Foam Behaviour in Liquid FeO-CaO-SiO2 Slags,” Scand. J. Metallurgy, 22:83 (1993) van Wie, B. J., Brouns, T. M., Elliott, M. L., and Davis, W. C., “A Novel Continuous Centrifugal Bioreactor for High-Density Cultivation of Mammalian and Microbial Cells,” Biotechnol. Bioeng. 38:1190 (1991) van der Jagt, M. H., Kleijntjens, R. H., van der Lans, R. G. J. M., and Luyben, K. C. A. M., “A Continuous Three Phase Suspension Reactor for the Microbial Decontamination of Excavated Polluted Soils,” 3rd Int. Conf. on Bioreactor and Bioprocess Fluid Dynamics, (A. W. Nienow, ed.), p. 335, Information Press, Ltd., Oxford (1993) van der Wielen, L. A. M., Potters, J. J. M., Straathof, A. J. J., and Luyben, K. C. A. M., “Integration of Bioconversion and Continuous Product Separation by Means of Countercurrent Adsorption,” Chem. Eng. Sci., 45:2397 (1990) Venkat, R. V., Stock, L. R., and Chalmers, J. J., “Study of Hydrodynamics in Microcarrier Culture Spinner Vessels: A Particle Tracking Velocimetry Approach,” Biotechnol. Bioeng., 49:456 (1996) Vorlop, K. D., Estape, D., and Gòdia, F., “Design of Ca-Alginate Immobilized Yeast Cell Beads with Controlled Low Density to Enhance Their Fluidization Behaviour in Bioreactors,” Biotechnol. Tech., 7:287 (1993)

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Webb, O. F., Davison, B. H., Scott, T. C., and Scott, C. D., “Design and Demonstration of an Immobilized-Cell Fluidized-Bed Reactor for the Efficient Production of Ethanol,” Appl. Biochem. Biotechnol., 51/52:559 (1995) Weng, D., Cheng,. L., Han, Y., Zhu, W., Xu, S., and Ouyang, F., “Continuous Ethanol Fermentation in a Three-Phase Magnetic Fluidized Bed Bioreactor,” AIChE Symp. Series, 88(289):107, AIChE, New York (1992) Weuster, D., Aivasidis, A., and Wandrey, C., “Ethanolfermentation of Sugar Containing Wastes with Zymomonas mobilis in a Fluidized Bed Reactor,” DECHEMA Biotechnol. Conf., 3:507 (1989) Weuster-Botz, D., “Continuous Ethanol Production by Zymomonas mobilis in a Fluidized Bed Reactor. Part I. Kinetic Studies of Immobilization in Macroporous Glass Beads,” Appl. Microbiol. Biotechnol., 39:679 (1993) Weuster-Botz, D., Aivasidis, A., and Wandrey, C., “Continuous Ethanol Production by Zymomonas mobilis in a Fluidized Bed Reactor. Part II. Process Development for the Fermentation of Hydrolysed B-Starch without Sterilization,” Appl. Microbiol. Biotechnol., 39:685 (1993) White, D. E., Gandek, T. P., Pikulin, M. A., and Friend, W. H., “Importance of Reactor Design in High-Consistency Ozone Bleaching,” Pulp & Paper Can., 94(9):16 (1993) Wisecarver, K. D., and Fan, L. S., “Biological Phenol Degradation in a GasLiquid-Solid Fluidized Bed Reactor,” Biotechnol. Bioeng., 33:1029 (1989) Worden, R. M., and Donaldson, T. L., “Dynamics of a Biological Fixed Film for Phenol Degradation in a Fluidized-Bed Bioreactor,” Biotechnol. Bioeng., 30:398 (1987) Wu, K. Y. A., and Wisecarver, K. D., “Biological Phenol Degradation in a Countercurrent Three-Phase Fluidized Bed Using a Novel Cell Immobilization Technique,” AIChE Symp. Series, 86(276):113 (1989) Yabannavar, V. M., and Wang, D. I. C., “Extractive Fermentation for Lactic Acid Production,” Biotechnol. Bioeng., 37:1095 (1991) Yang, S. T., “A Novel Gradient Particle Bed Bioreactor for Mixed Culture Fermentation,” presented at the AIChE Annual Meeting, New York (1987) Yeager, R., “New Profits from Advanced Control,” PIMA Magazine (Paper Industry Management Association), 77(1):6 (1995) Yee, C. J., Hsu, Y., and Shieh, W. K., “Effects of Microcarrier Pore Characteristics on Methanogenic Fluidized Bed Performance,” Wat. Res. 26:1119 (1992) Yongming, L., Yi, Q., and Jicui, H., “Research on the Characteristics of Start Up and Operation of Treating Brewery Wastewater with an AFB Reactor at Ambient Temperatures,” Wat. Sci. Tech., 28(7):187 (1993)

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682 Fluidization, Solids Handling, and Processing Zellner, G., Geveke, M., de Macario, E. C., and Diekmann, H., “Population Dynamics of Biofilm Development During Start-Up of a Butyrate-Degrading Fluidized-Bed Reactor,” Appl. Microbiol. Biotechnol., 36:404 (1991) Zellner, G., Neudörfer, F., and Diekmann, H., “Degradation of Lactate by an Anaerobic Mixed Culture in a Fluidized-Bed Reactor,” Wat. Res., 28:1337 (1994) Zhang, L., and Oeters, F., “A Model of Post-Combustion in Iron-Bath Reactors, Part 1: Theoretical Basis,” Steel Res., 62:95 (1991a) Zhang, L., and Oeters, F., “A Model of Post-Combustion in Iron-Bath Reactors, Part 2: Results for Combustion with Oxygen,” Steel Res., 62:107 (1991b) Zhang, L., and Oeters, F., “A Model of Post-Combustion in Iron-Bath Reactors, Part 3: Theoretical Basis for Post-Combustion with Preheated Air,” Steel Res., 64:542 (1993a) Zhang, L., and Oeters, F., “A Model of Post-Combustion in Iron-Bath Reactors, Part 4: Results for Post-Combustion with Preheated Air,” Steel Res., 64:588 (1993b)

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10 Dense Phase Conveying George E. Klinzing

1.0

INTRODUCTION

The term dense phase became popular in the conveying business in the mid 1970’s when a particular pneumatic conveying vendor coined the phrase to mean the pulsed piston type of flow that is often encountered in the handling of plastic pellets. Before that date, dense phase usually meant a high concentration of solids moving with a heavy concentrated layer at the bottom of the pipe. Thus one is faced with a dilemma of the exact meaning of dense phase conveying. Presently, one must use the words dense phase with other descriptors attached so as not to confuse and mislead people of the exact condition of flow explored. Figures 1 and 2 were prepared by Dhodapkar (1991) and present a comprehensive grouping of terms that could all generically be called dense phase conveying. Figure 3 shows some specific terms that have been used by particular vendors in their search to be different and to have a better and unique system. A comprehensive review of dense phase transport has been prepared by Konrad (1986) and can be referred to for further details.

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Figure 3. Classification of dense phase systems by vendors.

One can construct a number of different systems that would convey material in the dense phase condition. Figures 4–16 is a compilation of these systems indicating their unique characteristics. The Gatty system has air injected from a sub-pipe mounted inside the main conveying line. Since this system is without controls, it performs better for well-behaved plastic pellet systems. It is difficult to maintain a low velocity conveying condition with this system, Marcus et al. (1990). The Buhler Fluidstat system was first developed by Munschelknautz using a series of bypass pipes every half meter or so. Solids having good air retention properties are handled nicely with this system. A variation of this technique is the Moller Turbuflow system. The designers of this system claim that turbulence is regenerated in the conveying line. A trace air system uses a series of boosters employing a pressure regulator and a check valve. The Semco systems is depicted in Fig. 6. Dynamic Air incorporates a similar concept controlling the air volume rather than the pressure. Warren Springs Laboratories has developed a pulsed conveying system consisting of a flow vessel and an air knife downstream from the discharge valve as shown in Fig. 8. The plug lengths are controlled by the air knife operation time. An intermittent solids feed system shown in Fig. 9 can also control the size of the plug from both the air injection time and the discharge valve timing. Buhler has also developed a Takt-Schub system which incorporates a double pulse system. The blow vessel has two gas inlet valves whose timing can be controlled. The system is operated in an alternating pulsing manner.

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Dense Phase Conveying

687

Figure 4. Gatty’s dense phase system.

Secondary

Line

1-

"""/~

~

Figure 5. Turbuflow dense phase system.

Air

Figure 6. Trace air dense phase system (Semco).

11-Oct-2001 JMR


Figure 7. Dynamic Air dense phase system.

Figure 8. A pulse phase system arrangement.

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689

Figure 9. Intermittent solids feed system.

Figure 10. Takt-Schub system by Buhler.

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690 Fluidization, Solids Handling, and Processing Another system is the plug pulse system seen in Fig. 11. The discharge valve is alternatively activated with a gas pulse. This system works well for free flowing materials. The minipot is a variation of the plug pulse system. The operation of the minipot is depicted in Fig. 12. Waeschle (Fig. 13) has developed a pressure sensing plug system. The boosters in this system are actuated by pressure rises in the line. The Fluid-Schub system of Buhler, Fig. 14, uses valves to inject air from an auxiliary air line. A system that produces extruded flow of long plugs can also be devised as shown in Fig. 15. A new low velocity conveying system has been devised by Molerus and Siebenhaar (1990) and is shown in Fig. 16. This system relies on vibrations being put into the system by an unbalanced motor. These vibrations cause the materials to move with a low velocity throughout the system. This system is particularly useful for fragile materials. In viewing all these systems one is faced with the over riding question on how to analyze, predict and scale such operation. For some designs, unfortunately, the only solution is to build a unit and test it. Others have some basic physics that can be reviewed. Assumptions for analysis can be made which often result in idealized systems rather than the real thing. This condition is not all bad because now one has a limit or a bound that can bracket the design and scale operations.

\

Air

~

Figure 11. Plug pulse conveying system.

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691

Air

Conveying

Air

Filling

Figure 12. Minipot conveying system.

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693

Figure 16. Molerus dense phase transport system.

2.0

ADVANTAGES OF DENSE PHASE CONVEYING

There have been a number of advantages that have been listed for dense phase conveying by Paul (1990). These are quoted as (i) Decreased energy usage due to significantly reduced volumes of air. (ii) Reduced material breakage or degradation due to slower conveying velocities. (iii) Reduced pipeline wear due to lower conveying velocities. (iv) Smaller conveying pipeline sizes due to heavier line loading capabilities. (v) Smaller dust collection requirements at the material destination due to lower conveying air volumes. In addition Paul has prepared a chart listing the characteristics of the dense phase conveying in comparison to dilute and medium dense phase systems as shown in Fig. 17. Vendors have used certain types of dense phase conveying operations. In the trade, these vendors and their unique modes of transport have been classified in Fig. 3.

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11-Oct-2001 JMR Figure 17. Pneumatic conveying system classification (by Paul, 1991).

Dense Phase Conveying 3.0

695

BASIC PHYSICS

A starting model can be selected depending on the type of dense phase transport. Modeling a homogeneous dense phase would use the same approach as dilute phase with a new frictional term. This approach would have two contributions due to the gas alone and the linear combination with the solids contribution. Eq. (1)

∆P = ∆Pgas + ∆Psolids

Mathur and Klinzing (1981) have developed a frictional term to account for the solids pressure drop which is meant for high loading systems (µL= 10 to 50). Their expression for f s is Eq. (2)

f s = 55.5D

1.1

(U

0. 64 0.26 0.91 g dp ρp

)

where the pressure drop for the solid contribution is Eq. (3)

(∆P L)solids = f s (1 − ε ) ρ pU 2f

D

A work by Stegmaier (1978) has also developed a friction factor associated with the solids flow. This expression is Eq. (4)

−0. 3

−1

f s = ( 2.1/ 4)µ L Fr Fr

*0.25

(d p / D)

−0.1

where Eq. (5)

( ∆P / L ) solids = µ L f s ρU f / D 2

and Eq. (6)

Fr = U f / gD

Eq. (7)

Fr = U t / gD

2

*

2

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696 Fluidization, Solids Handling, and Processing Weber (1973) has developed a simple expression for dense phase which can be written as − dP / P = βgµ L dl /( RTVp / V f )

Eq. (8)

This equation can be integrated to yield

(

where

[

P1 = P2 exp βgµ Ll R′TVp / V f

Eq. (9)

])

β = sinα + µRcosα µR = coefficient of sliding friction

Wen and Simon (1959) have recommended an overall expression that lumps the air and solids pressure contributions together. ∆P / L = 3. 79W (144 *Vg 0 .55 )(d p / D)

Eq. (10) where

0. 25

W - lb/ft2s Vg - ft/s

The Wen and Simon expression has often been used to calculate the pressure loss in systems where one has a high concentration material at the bottom of the pipeline. This type of flow is often called two-phase flow. Wirth and Molerus (1986) have termed this two-phase flow as strand flow and have developed a technique to predict the pressure loss in such flows. Using Figs. 18 and 19 one can follow the procedure: (i) Calculate the friction number Fri from ρp 2 2 Fri = V g Eq. (11) [ ρ (1 − ρ / ρ p )(1 − ε ) Dgfr]

where

ε = voidage of strand (0.4) fr = coefficient of sliding friction = 0.6

(ii) The loading µl is set and ρ µ/[ρ p(1 - ε)] is calculated. (iii) These values locate a point on the Fig. 18 and from the ordinate the reduced pressure drop can be found.

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697

With a simulation technique Peng et al. (1994) have studied the particle-particle collisions with rotational and floating mechanisms of coarse particles in horizontal flow for higher concentrations of particles. Figure 20 gives the distribution of particles across the pipe cross-section. One notes that higher concentrations are seen at the bottom of the pipe. 0.3

~

0.2

~ -;:; t -:-.

0-IQ. t ~

0.1

~

--, I

./ II

~

0.06

;;'1 I ,

0

.0

FIJ= I..,p~

v .v.

.

-,~-

Figure 18. Additional pressure drop and state diagram for stable conveying particle with sliding strands (ε = 0.4). (Wirth and Molerus, 1986.)

11-Oct-2001 JMR


1.3 10..

~ > .

8 ~

..;.;.

,

6

0.02 \

~

Fiz

~

2x104

0

200

100

.Jp,/p

300

360

Fr.

Figure 19. State diagram for suspended flow conveying in horizontal pipes. (Wirth and Molerus, 1986.)

:j

~~I 20

..1

Figure 20. Dense phase simulation by Peng, Tomita and Tashiro (1991).

4.0

PULSED PISTON FLOWS

Welschof (1962) carried out tests on dense phase plugs having a low velocity of 1 m/s. Later Lippert (1965) did an systematic analysis of the plug flow behavior. Following these researchers in the plug field were Weber (1973), Konrad et al. (1980) and Legel and Schwedes (1984). In

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699

these analyses, a moving packed bed model was proposed using the theory of soil mechanics. An experimental study by Dickson et al. (1978) showed how a plug moved by mechanical means requires a force that increases exponentially with plug length. A simple force balance for a cohesive powder inclined in a plug at an angle of α can be given as dP

Eq. (12)

dZ

=

d σ Ax dZ

+

4τ w D

+ g sin α

Konrad was the first to address the issue of pulsed piston transport using the properties of the solids as they slide through the pipe in a plug-like motion. The friction generated in such systems often can be likened to bin and hopper flow and design, requiring shear stress measurements such as carried out by the Jenike shear stress unit. The final expression using the Konrad approach can be written for horizontal flow as Eq. (13) ∆P Lp

=

4µ w K w F ′ D

+

4µw ( K w + 1) c′ cosφ cos(ω − φ w ) D

+ 2ρ b g tanφ w + 4Cw / D

The stress terms F´ at the front of the plug and B´ at the back of the plug, dependent on the pressure drop and powder properties, can be developed from a momentum balance but often times they are set equal to each other. Using the momentum balance

Eq. (14)

F'= ρb

dVp

2

1− d

where Eq. (15)

d=

1 1 + c ' / 0 .542 ( gD)

1/ 2

The term c´ is zero for cohesionless powders such as plastic pellets. As one sees, there are several parameters that need to be measured specifically

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700 Fluidization, Solids Handling, and Processing for the particular material conveying. This type of data is difficult to generalize which makes pulse piston modeling particularly difficult. Other investigators have studied the pulse piston flow operation (Legel and Swedes, Lippert, Aziz and Klinzing, 1988). Pulse piston operations are often demonstrated at trade shows where plastic pellets are used to demonstrate the concept. Plastic pellets are unique materials in that they can actually form their own pulsed pistons without special feeding devices. Cohesive materials require special feeder arrangements as has been seen in the specific examples shown above. Weber (1973) has put forth a rather simple approach to pulsed piston modeling which couples the gas and solid behavior. The expression that Weber suggested is given in Eq. (9). One notes that this expression has an exponential decay of the pressure as the plug length decreases. This behavior has often been seen in industrial operations. A series of studies were performed by Borzone and Klinzing (1987), Gu and Klinzing (1989) and Aziz and Klinzing (1990) on plug flow of coal. In order to transport plugs of fine materials, especially in the vertical case, it is necessary to have an sudden application of pressure so that the transport gas does not leak through the plug. It is very apparent that there exists a minimum plug length for material transport in this mode or else there will be a breakthrough of the gas leaving a mound of material on the bottom of the pipe. Their modeling was confined to modifying the work of Konrad for cohesive flow, as well as that of Weber. Good agreement was seen between the experiment and model as shown in Fig. 21. When increased pipe diameters were studied in vertical flow, the minimum plug length did not vary significantly with pipe diameter while the maximum plug lengths decreased with pipe diameter to maintain stable plugs. The vertical plugs were dominated by the gravitational force. The concept of plug length and optimum operation is one that often is of concern. Aziz, in his study of fine coal, found that there is a minimum plug length that will form in order to convey in the piston model. This minimum condition is of interest in the design of any pulse piston operation. Unfortunately, the ability to predict such a minimum length is lacking. Known powder properties are essential to predict the minimum plug lengths. Figure 22 shows the various modes of collapse of the horizontal plugs when the plug length is below the minimum value. It has also been observed by Aziz that a layer of material can be deposited on the bottom of the pipe and a plug-like flow will move with this as its base, thus transporting in a half-plug mode. The pressure drop across a plug can be represented by a straight line behavior as predicted by the

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701

Konrad model for relatively short plug. Longer plugs (3 meters +) show an exponential pressure drop across the plug. Figure 23 indicates this behavior observed by Aziz. This nonlinear increase in pressure across a long plug can be observed in a simple experiment developed by Arnold (1992). A plug of solids supported by a loose fitting disk attached by a wire over a pulley shows that as the plug length increases the force necessary to move the plug increases exponentially. Addition of a small amount of lubrication or fluidization air shows a marked decrease in the force to move the plug for the apparatus shown in Fig. 24.

Figure 21. Pressure drop versus plug length. (Aziz and Klinzing, 1990.)

I

""

1

2

3

Figure 22. Hypothesis of plug collapse (increasing gas leakage with position).

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702 Fluidization, Solids Handling, and Processing PluQ Velocity = 0.2 m/s

.. Do .x 0: ~ C .. ~ .. .. .. 11:

0

0.1

0.2

0.3

0.4

05

06

Plug

Length,

m

0.7

0.8

Figure Figure23. 23. Pressure Pressure drop drop for for Pittsburgh Pittsburghcoal coalfor foraadiscrete discreteplug plugflow. flow.

1*1

Scale

Force Granular Material

Metal -Air

Figure 24. Arnold’s plug flow device.

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Disc Vent

0.9

1


703

Recently, Destoop (1994) has addressed a system consisting of a piston being conveyed in a pulse piston fashion. Likewise Plasynski et al. (1994) have also addressed this type of flow giving an expression for modeling and predicting he minimum plug length. The minimum plug length can be found by solution of a force balance shown in Fig. 25 and represented as Eq. (16) where

AT dP = σAz AT + τwAc1 +τAi AT = transverse cross sectional area at failure = (1/8)D2(θ - sinθ) Ac1 = circumferential area at failure Ai = interfacial powder area at failure = D sin(θ /2)Lsp

In addition, for the system one sees as the line length increases the length of the gas plugs will expand. Figure 26 shows an example of this condition for a specific example.

Figure 25. Plug flow shear mechanism.

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Figure 26. Plug and air gap flow behavior with distance.

Recently Min (1994) under the direction of Wypych has experimentally studied the conveying of plastic particles and wheat over a sizable distance in the pulsed piston operation. The results that they obtained experimentally agree qualitatively with the model suggested by Plasynski et al. and Destoop. Figure 27 exhibits this length of the air gap as a function of air or solids flow rates at constant air or solids mass flow. There is design strategy for this type of flow that can be suggested. Tsuji et al. (1990) have modeled the flow of plastic pellets in the plug mode with discrete dynamics following the behavior of each particle. The use of a dash pot/spring arrangement to account for the friction was employed. Their results show remarkable agreement with the actual behavior of real systems. Figure 28 shows these flow patterns. Using models to account for turbulent gas-solid mixtures, Sinclair (1994) has developed a technique that could have promise for the dense phase transport. Further analysis of plug flow has been given by Destoop and Russell (1995) with a simulated computer model for catalyst and polymer materials. The model was developed based on piston-like flow of plugs separated by plugs of gas. The model has been employed taking into account the product grade, temperature, flow rates and line configuration.

11-Oct-2001 JMR

0 (\I .-

..;..;00 ."X.

O 00

:Q~ -o

~lI

w 'de~

o v

0

0

o d

C\I

0

o.;r o

0

o

,~

co 0 d

Ta


J!'V' JO 4~6ual

705

11-Oct-2001 JMR

Figure 27. Air flow gap length versus mass flow rate of air for plastic pellets. (Min, 1994.)


Figure 28. Plug flow simulation. (By Tsuji et al., 1990.)

5.0

VERTICAL FLOW SYSTEMS

The use of recirculating fluid beds has caused considerable interest in dense phase vertical conveying. These units are indeed dense phase transport systems with a significant amount of recirculation taking place.

11-Oct-2001 JMR


707

There are several researchers who have addressed this operation. Of particular note is the modeling of Yang (1988) and the analysis of Knowlton (1994) for particularly large diameter units. Yang presented a model that describes the dynamics of a circulating fast fluidized bed. The dilute phase region was correlated with a set of equations he previously developed for this type of flow. The pressure drop in the dense phase with an annular region is given as Eq. (17)

∆P = ρ p (1 − ε ave ) L

with the solids inventory in this region as Eq. (18)

I = ρ p (1 − ε ave ) Ar L

The height of the dense phase L is obtained by a pressure balance around the complete circulating fluidized bed loop. Good agreement is seen with this model and the existing data in the field. Knowlton has cautioned on the difference between small diameter and large diameter systems for pressure losses. The difference between these systems is especially apparent for dense phase flow where recirculation occurs and wall friction differs considerably. Li and Kwauk (1989, 1989) have also studied the dense phase vertical transport in their analysis and approach to recirculating fluid beds. Li and Kwauk’s analysis included the dynamics of a vertical pneumatic moving bed upward transport using the basic solid mechanics formulation. Some noncircular geometries were treated including experimental verification. The flows have been characterized into packed and transition flows. Accurate prediction of the discharge rates from these systems has been obtained. Vertical pneumatic moving bed transport was also explored with a theoretical and experimental study by Li and Kwauk (1989). A criterion was put forth for stable operations which include absence of bubbles, slugs and voidage pockets, and having a constant mean particle velocity. There should be a means of suppressing fluidization and bubbling and sufficient hydrodynamic drag to permit lower-layer particles to be lifted to the upper layers. With a pressure drop of 1.5 times that to overcome gravity, the critical pneumatic moving-bed transport can occur.

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708 Fluidization, Solids Handling, and Processing Li (1994) has also studied the mechanics of arching in moving-bed standpipe flow. He was able, for this downflow situation, to obtain the critical arching span which agrees with reported data. The critical or minimum radius Rc for no arching is given as

Eq. (19)

6.0

Rc =

f c sin 2 β '

1 (1 + sin δ ) sin 2 β ' [ ρ b g + ∆P / ∆h ] [1 − k ] (1 + cos 2 β sin δ ) 2 tanδ

BOOSTERS

The topic of boosters is often seen as either the salvation or demon of dense phase transport. The design of the appropriate placement of booster is still evasive. Often too many boosters are utilized such that the overall gas usage turns out to have a dilute phase system. The Waeschle Company has designed a unique booster system that has individual controls. This design can handle most situations; for example, if a long plug forms by the union of two plugs or other flow hindrances, the system responds by injecting air to assist in the transport. This action breaks the plug into more manageable plug lengths. Overall care must be taken in using the boosters. Boosters can help in the conveying of a variety of difficult materials. When the flow causes stresses to be built up at bends and other stress points, the proper application of a booster can provide the added air to permit troublefree conveyance. One needs to remember that a particular booster application that worked for one material will probably not work for another. Secondary air injection for the transport of solids has seen little analysis in the literature despite the prevalence of this technology in industry. One exception is the recent work of Dhodapkar et al. (1995). This analysis employs the use of a sizable pilot unit for developing operational guidelines for Geldart type B and D, free flowing, plastic pellets and granular polymers. Booster air was shown to lower the average velocity, permit higher operating pressures, and reduce the unstable operation of the system. Pressure fluctuations were significantly reduced for the Geldart type B material when using boosters. They also found that flow can be initiated successfully in a full line at lower conveying pressures with booster air injection.

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709

NOTATIONS A C c´ cw D dp F fs Fr g K Kw l L n P R T Ug Uf V W z

-

Area Internal cohesion coefficient Cohesion coefficient Wall cohesion coefficient Diameter of pipe Particle diameter Stress at the front of the plug Solids friction factor Froude Number Gravity constant Coefficient of internal friction (1-sin 2f)/(1+sin 2f) Coefficient of internal friction at the wall Length of sector at failure Length of plug or transport distance Number of plugs in the pipeline Pressure Gas law constant Temperature Superficial velocity Fluid velocity Volume Solids flux Distance

Greek α - Angle of inclination ε - Voidage φ - Angle of internal friction φw - Angle of wall friction θ - Sector angle at failure ρ - Gas density ρb - Bulk density ρp - Particle density σ - Normal stress τ - Shear stress τw - Wall shear stress µL - Solids loading sinω ~ sinφ w/sinφ

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710 Fluidization, Solids Handling, and Processing Subscripts C1 i sp T w z

-

section at failure integer, interfacial plug total wall distance

REFERENCES Arnold, P., Personal communication (1992) Aziz, Z. B., and Klinzing, G. E., “Plug Flow Transport of Cohesive Coal: Horizontal and Inclined Flows,” Powder Tech., 55:97–105 (1988) Aziz, Z. B., and Klinzing, G. E., “Dense Phase Plug Flow Transfer: The 1-inch Horizontal Flow,” Powder Tech., 62:41–49 (1990) Borzone, L. A., and Klinzing, G. E., “Dense Phase Transport: Vertical Plug Flow,” Powder Tech., 53:273 (1987) Destoop, T., “Mathematical of Dense Phase Conveying Systems - A Revolutionary Move Towards ‘Smart’ System Control,” Proc. of 1994 Bulk Solids Handling Conf., Rosemont, IL (1994) Destoop, T., and Russell, B., Proceedings of the Fluidization and Particle Systems Sessions of AIChE Annual Meeting, pp. 351–355, Miami, FL (1995) Dhodapkar, S. V., Jacob, K., and Merrick, D., Proc. of the Fluidization and Particle Systems Sessions of AIChE Annual Meeting, Miami, FL (1995) Dhodapkar, S. V., Personal communications (1991) Gu, H., and Klinzing, G. E., “Vertical Plug Flow of Cohesive Coal in 2- and 4Inch Pipes,” Powder Tech., 57:59-67 (1989) Knowlton, T., “The Effect of the System Pressure/Pipe-Diameter/Mass Flux Interaction on Pressure Drop in Dilute Phase Pneumatic Conveying,” Proc. of Pneumatic Conveying Workshop, Powder Technol. Forum, AIChE, Denver (1994) Konrad, K., “Dense Phase Pneumatic Conveying: A Review,” Powder Tech., 49:1 (1986) Konrad, K., Harrison, D., Nedderman, R. M., and Davidson, J. F., Proc. of Pneumotransport 5, Paper E1, BHRA, p. 225, Cranfield, Bedford, UK (1980)

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711

Legel, D., and Swedes, J., “Investigation of Pneumatic Conveying of Plugs of Cohesionless Bulk Solids in Horizontal Pipes,” Bulk Solids Handling, 4(2):53– 59 (1984) Li, H., “Mechanics of Arching in a Moving-Bed Standpipe with Interstitial Gas Flow,” Powder Tech., 78:179–18 (1994) Li, H., and Kwauk, M., “Vertical Pneumatic Moving-Bed Transport - I. Analysis of Flow Dynamics,” Chem. Engr. Sci., 44:249–259 (1989) Li, H., and Kwauk, M., “Vertical Pneumatic Moving-Bed Transport - II. Experimental Findings,” Chem. Engr. Sci., 44:261–271 (1989) Lippert, A. “Die Staub-Luft-Forderung Pulvern und Schuttgutern mit hohen Konzentrationen im Gasstrom-ein neuer Fordergang,” Dissertation TH. Univ. of Karlsruhe (1965) Marcus, R. D., Rizk, F., Klinzing, G. E., and Leung, M., “Pneumatic Conveying of Solids,” Chapman Hall (1990) Mathur, M. P., and Klinzing, G. E., “The Dense and Extrusion Flow Regime in GasSolid Transport,” Can. J. of Chem. Engr., 59:590–594 (1981) Molerus, O., and Siebenhaar, W., “Vibration Induced Pneumatic Conveying,” Pneumatech IV, p. 213, Glasgow, Scotland (1990) Min, B., “Ph.D. Dissertation, University of Wollongong, Australia (1994) Paul, K. D., “Dense Phase Pneumatic Conveying Improving Efficiency,” Powder and Bulk Engr., p. 41 (1991) Peng, X., Tomita, Y., and Tashiro, H., “Effect of Particle-Particle Collision and Particle Rotation upon Floating Mechanism of Coarse Particles in Horizontal Pneumatic Pipe,” JSME Intl. J., Series B., 37(3):485–490 (1994) Plasynski, S. I., Dhodapkar, S. V., and Klinzing, G. E., Powder Tech., 81:3–7 (1994) Sinclair, J., Proc. of Pneumatic Conveying Workshop, Powder Technology Forum, AIChE., Denver (1994) Stegmaier, W., Fordern und Heben, 28:363 (1978) Tsuji, Y., Tanaka, T., and Ishida, T., “Graphic Simulation of Plug Conveying,” Proc. of Pneumatech, 4:39–50 (1990) Weber, M., ”Stromungs-Foerdertechnik,” Krausskopf Verlag, Mainz (1973) Welschof, G., ”Pneumatische Forderung bei Grossen Fordergutkonzentrationen,” VDI-Forschungscheff, p. 492, VDI Verlag, Dusseldorf (1962) Wen, C. Y., and Simons, H. P., AIChE J., 6:263–8 (1959) Wirth, K. E., and Molerus, O., “Critical Velocity and Pressure Drop for Horizontal Flow,” 9:156–166 (1986) Yang, W. C., “A Model for the Dynamics of a Circulating Fluidized Bed Loop,” Second Int. Conf. on Circulating Fluidized Bed, Compiegne, France (1988)

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11 Design Considerations of Long-Distance Pneumatic Transport and Pipe Branching Peter W. Wypych

1.0

INTRODUCTION

The pneumatic transportation of bulk solids through pipelines has been in existence for over one hundred years. Some of the earlier applications included the vacuum unloading of grain from ships and the extraction of sawdust from timber mills. This method of conveying now is being selected for an increasing number of industrial applications and products and hence, is playing a more vital and integral role in numerous bulk handling operations and processes. Some of the major reasons are listed below. • Enclosed, safe and environmentally attractive method of transport suitable to a wide variety of products, including those with bacteria-prone, toxic or explosive properties.

712

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Long Distance Pneumatic Transport and Pipe Branching

713

• Simple systems requiring a prime mover, a feeding device, a conveying pipeline and a cleaning or disengaging device (Marcus et al., 1990). • Flexibility in pipeline layout. • Ability to distribute product to a number of different areas within a plant and/or pick up material from several different locations. • Low maintenance and manpower costs. • Multiple use—a single pipeline can be used for a variety of products. • Ease of automation and control. Unfortunately, many new and old systems have not been designed properly and are being operated inefficiently. Some of the major consequences include high energy consumption, excessive system erosion, inadequate conveying capacity, unexpected pipeline blockages, excessive product damage and hence, poor quality control and/or increased maintenance. These problems have resulted mainly from • A lack of appreciation of product properties and/or characteristics • Inadequate and/or inaccurate design procedures • Selecting unsuitable hardware and/or mode of transport (Wypych, 1995a) • Improper commissioning and/or operation of the plant A significant number of developments have occurred over the past decade to address these important issues of pneumatic conveying (Wypych, 1995a). This chapter summarizes some of the major design considerations that have resulted from this work in relation to long-distance and pipe branching applications.

2.0

LONG-DISTANCE PNEUMATIC CONVEYING

With the advent of high-pressure feeders after World War II (e.g., blow tanks, screw pumps), many designers and users began to apply pneumatic conveying to much greater distances and/or capacities (e.g.,

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714 Fluidization, Solids Handling, and Processing L ≥ 1 km and/or ms ≥ 100 th-1). However, in many cases, this resulted in various problems, such as • higher than expected pressures and energy consumption, • excessive transport velocities, system erosion (e.g., pipes, bends) and product damage, • low system reliability (mainly due to premature failure of valves and instrumentation), • inadequate conveying capacity, and/or • unexpected material buildup and/or blockages along the pipeline. These problems in turn were hampering the successful design and/or operation of long-distance pneumatic conveying systems and hence, the future potential of this method of transport to a wide variety of industries. To overcome these problems, a number of interesting developments in research and technology have occurred over recent years and these have resulted in the following important design considerations. • Characterization and classification of the material(s) to be conveyed. • Improved blow tank design to ensure an efficient and controlled discharge rate of material. • New test-design procedures for the accurate prediction of pipeline pressure drop, including the effects due to horizontal/ vertical flow and bends. • Stepped-diameter pipelines to minimize pressure drop, velocity, wear and power consumption. • Reliable valves for blow tank filling, venting and discharge. • Manual or automatic back-pressure unblocking of conveying pipelines. 2.1

Product Characterization and Classification

The characterization and classification of bulk solids is becoming an increasingly important design requirement to assess the suitability of

11-Oct-2001 JMR


715

pneumatically conveying a material over long distances. The current trends and experiences indicate that the major influential parameters include • Particle size and distribution • Particle density, bulk density and particle shape • Cohesive, wet, sticky and/or electrostatic properties (if applicable) • Temperature of product and carrier gas • Permeability (or fluidization) and deaeration characteristics of the material (Mainwaring and Reed, 1987; Jones and Mills, 1989), which depend strongly on the above parameters Possibly, the most difficult aspect of determining particle size is selecting initially the correct or relevant definition and then calculating a mean or average diameter to represent the complete bulk solid. To some extent, this will depend on the following. • The measuring apparatus and its principle of operation • The final application or requirements (e.g., determination of free-settling velocity v∞ , minimum fluidization velocity Vmf , minimum transport velocity Vfmin , and/or pipeline air pressure drop ∆p t) • The basis of definition used in a theoretical or empirical relationship (e.g., sieve or volume measurement) In some cases, especially for very fine powders, researchers have looked at other properties to explain or classify product behavior. For example, Geldart et al. (1984) have found that the ratio of “tapped” to “aerated” bulk density provides a good indication of the likely fluidization characteristics of fine and cohesive powders. However, this section pursues particle size measurement and evaluates its importance (as well as density) for the purpose of classifying the suitability of powders for long-distance pneumatic conveying applications. Initially, an appreciation of the fundamentals and the existing powder classification techniques is required. Physical Properties. To determine or calculate particle size, it is important to be aware of the following different definitions and related properties (Geldart and Abrahamsen, 1981; Allen, 1975).

11-Oct-2001 JMR

716 Fluidization, Solids Handling, and Processing d p = Arithmetic mean of adjacent sieve sizes. d pm = Mean particle size from a standard sieve analysis,

= Eq. (1)

Σ (∆ M )  ∆M   Σ  d   p 

where ∆M is the mass percent of product between adjacent sieves. d pwm = Weighted mean diameter (Allen, 1975) based on sieve analysis,

(

Σ ∆M d p = Eq. (2)

Σ(∆M )

)

d sv = Diameter of a sphere with the same surface area to volume ratio as the particle. d svm = Mean surface volume diameter,

= Eq. (3)

Σ (∆ M )  ∆M   Σ  d   sv 

d v = Diameter of a sphere with the same volume as the particle. d vm = Mean (equivalent) volume diameter,

= Eq. (4)

Σ (∆ M )  ∆M   Σ  d   v 

d vwm = Volume weighted mean diameter (Allen, 1975),

Σ (∆ M d v ) = Eq. (5)

11-Oct-2001 JMR

Σ (∆ M )


717

d 50 = Median particle diameter (Allen, 1975). d v50 = d v50 for a volume diameter distribution. d v50 = d v50 for a sieve size distribution. Ψ = Particle sphericity (Geldart and Abrahamsen, 1981). = Eq. (6)

d sv dv

Assuming the following results of Geldart and Abrahamsen (1981) and knowing (or assuming) appropriate values of sphericity, the various diameters d svm, d v50 , d p50 , and d vm may be determined for any given material. This information is particularly useful when it becomes necessary to compare the classification of different bulk solids, whose size distributions were determined on different machines or using different techniques (e.g., sieve analysis, volume-based measurement). dv 1.1 < d t-

'!

ASH

(GRITS)

POLYETHYLENE PELLETS

In z L.J

0

10

GRANULATEO SUGAR

10

11

L.J ..J u

PEARLITE

'2

FLOUR

13

PULUERISEO FUEL

ta: < 0..

ASH

0 Z < a: 0 tu < l1o Z 0

t< a: L.J < I L.J 0

"' ,

~

0

40

80

120

160

200

~-o PRESSURE DROP PER UNIT LENGTH , (mbor/m)

( ~

) mf

Figure 6. Pressure drop per unit length versus deaeration and particle density factor (Mainwaring and Reed, 1987).

Figure 5 is a plot of permeability factor κp versus (∆p/∆h)mf and contains the boundary Vmf = 50 mm s-1. Figure 6 is a plot of κd /ρ s versus (∆p/∆h)mf and contains the boundary X = 0.001 m3 s kg-1. By using both Figs. 5 and 6, it is possible to propose the following three criteria relating to the potential modes of pneumatic conveying in conventional pipelines.

11-Oct-2001 JMR


729

• If for a given material, Vmf > 50 mm s-1 (i.e., above the boundary shown in Fig. 5) and X < 0.001 m3 s kg-1 (i.e., below the boundary shown in Fig. 6), then dense-phase low-velocity slug-flow (Wypych, 1995a) is possible (e.g., mustard seed, polyethylene powder, 1000 µm sand, polyethylene pellets and granulated sugar). Note that dilute-phase also is possible. • If Vmf < 50 mm s-1 and X > 0.001 m3 s kg-1, then fluidized dense-phase (Wypych, 1995a) is possible (e.g., cement, pulverized coal, flour, fly ash). Note that dilute-phase also is possible. • If Vmf < 50 mm s-1 and X < 0.001 m3 s kg-1, then “fluidized” dense-phase or low-velocity slugging flow is not possible in a conventional pipeline (e.g., slate dust, zircon sand, fly ash grits). That is, single-slug conveying (Wypych, 1995a), bypass conveying (Wypych, 1995a) or dilute-phase may have to be considered for these materials. Application to Long-Distance Pneumatic Conveying. From the above three classifications, there is sufficient evidence to suggest that powder classification (i.e., to select ultimately the most suitable mode of conveying for a given product and its behavioral properties) depends on the following properties: • Particle size distribution and density. • Particle shape or sphericity (as indicated by the definitions of diameter). • Deaeration and permeability. • The ratio of tapped to poured (or perhaps fluidized) bulk density. • Diameter of conveying pipeline or particle to pipe diameter ratio. Also, it seems that most of these properties are interdependent. For example, deaeration and permeability (Mainwaring and Reed, 1987) and perhaps the bulk density ratio (Jones and Mills, 1989) seem to provide an adequate mechanism to detect changes in material performance due to different particle size distribution, density and/or shape. However, possibly the greatest disadvantage or limitation of these empirical techniques is the need to standardize the experimental apparatus and techniques. For exam-

11-Oct-2001 JMR

730 Fluidization, Solids Handling, and Processing ple, the measured values of deaeration rate depend on the size of the plenum chamber and to some extent the type of gas distributor. Also, different devices and techniques are available to determine vibrated or tapped bulk density. Standardization is necessary so that the results will be applicable on an international level and can be used/compared by other researchers. Based on the previous classifications and discussions, as well as the author’s own experiences, it is suggested that the following design procedure and considerations be adopted to provide an initial indication of densephase suitability, which as described below also has been found useful in assessing conveying performance over long distances. • For a given bulk solid, determine particle size distribution, median particle diameter d v50 (e.g., using a Coulter Counter or a Malvern Laser Diffraction Analyser) and ρ s. • Classify the bulk solid according to the Geldart (1973) fluidization diagram. This information is useful in estimating say, potential rat-holing problems inside blow tanks. If the material is relatively close to a classification boundary, then expect fluidization behavior from either one of the adjoining groups. • Test the bulk solid in a fluidization chamber to confirm both the Geldart (1973) classification and the material’s air retention properties (i.e., by undertaking deaeration experiments). • Classify the bulk solid according to the Dixon (1979, 1981) slugging diagram that matches the existing or proposed diameter of pipeline. Based on the research and consulting experience of the author, the following possibilities are proposed in relation to the Dixon (1979) classifications. — Typical Group A materials (e.g., cement, baghouse fly ash, pulverized coal, carbon fines) are the best candidates for “fluidized“ dense-phase (Wypych, 1995a) and long-distance pneumatic conveying. It is possible to achieve high values of solids/gas loading (e.g., m* = 150 over 200 m; m* = 25 over 1 km). These materials retain their aeration for considerable lengths of time.

11-Oct-2001 JMR


731

— Typical Group B materials (e.g., alumina, PVC powder, fine sand, castor sugar) may cause serious problems in fluidized dense-phase (e.g., unstable plugging, severe pipe vibrations, high pressures) and hence, may need to be transported in dilute-phase over long distances (i.e., using conventional pipelines). It still may be possible to convey these materials in “dense-phase” using low-velocity slugflow or bypass pipelines (Wypych, 1995a). Also, for the coarser Group B materials (e.g., fine sand, castor sugar), it may be possible to consider the alternative of single-slug dense-phase conveying (Wypych, 1995a). Of these three options, bypass conveying is preferred when long distances are involved, especially when L > 500 m. — Typical Group C materials (e.g., precipitator fly ash, lead fume, zinc dust) may behave like Group A powders (i.e., good fluidized dense-phase or longdistance performance). However, these materials can be quite cohesive and it is important to ensure buildup problems do not occur inside the blow tank feeder or along the pipeline. For example, precipitator fly ash with d v50 ≈ 10 µm has been conveyed successfully and efficiently over long distances, whereas lead fume with d v50 ≈ 5 µm has been found to cause buildup and eventual blockage problems inside the pipeline. Flexible hoses or “collapsible” pipelines may be used to prevent this buildup of material. Another option is to consider low-velocity plugging using either an air-knife at the beginning of the pipeline or a bypass conveying pipeline. In some cases, the plugs tend to be “self-cleaning”. However, these options do not become practical over long distances and must be considered carefully in terms of capital and maintenance costs. — Many typical Group D products (e.g., wheat, rice, sugar, plastic pellets, cereals, barley, malt, agglomerated milk powder) have relatively narrow

11-Oct-2001 JMR

732 Fluidization, Solids Handling, and Processing size distributions, high permeability and display natural slugging ability (Wypych and Hauser, 1990) and good low-velocity slug-flow (Wypych, 1995a). However, this mode of flow usually is limited to conveying distances of L < 500 m. The Group D materials that possess wide size distributions (especially a considerable amount of fines) and/or unusual particle shape (e.g., crushed coal, petroleum coke, crushed bath, rice hulls) usually are not suited to low-velocity conveying. In these cases, the options are dilute-phase or single-slug dense-phase (i.e., these materials usually are too coarse for bypass conveying). However, when long distances are involved, dilute-phase may be the only practical option. — If any material is in the vicinity of a classification boundary (i.e., Group A-B or Group B-D boundary), then due to particle size distribution it is possible that the material may exhibit flow behavior or performance from either one of the adjoining categories. Case Study. The above approach was applied recently to investigations into blockage problems occurring on a long-distance pneumatic conveying system handling baghouse fly ash at a coal-fired power station. It was reported by the operator that the blockages • occurred mainly at the beginning of the pipeline, • were accompanied by severe pipe vibrations, • eventually cleared themselves after a period of time. During a site inspection it was found that in between baghouse cleaning cycles, each blow tank was allowed to be filled with the coarser particles falling through the dust collection hopper. A sample of material was collected directly from a blockage inside the pipeline and also from the receiving silo which contained the conveyed product. The physical properties of the “successful” (fine) material were found to be d v50 ≈ 15 µm, ρ s ≈ 2200 kg m-3 and ρ bl ≈ 900 kg m-3, and for the “troublesome” (coarse) material d v50 ≈ 100 µm, ρ s ≈ 1900 kg m-3 and ρ bl ≈ 800 kg m-3. Note a considerable amount of unburnt particles was observed in the latter sample

11-Oct-2001 JMR


733

and this explained the somewhat lower value of ρ s. As a result of applying the powder classification techniques (Geldart, 1973; Dixon, 1981) and the other suggestions made in this chapter, it was found that • the fine ash was classified as a Group C powder—this material was conveyed successfully in fluidized dense-phase over long distances, • the coarse ash was classified as a Geldart (1973) Group A material—subsequent fluidization tests confirmed good fluidization characteristics but poor air retention properties (i.e., the material deaerated quickly as soon as the fluidizing air was removed), • the coarse ash according to Dixon (1981) was classified as Group B in a 50 mm NB pipe system and was located on the Group A-B boundary in a 100 mm NB pipe system (i.e., refer to Figs. 3 and 4). Based on these findings as well as the nature of the blockages, the coarse ash was considered as a typical Dixon (1981) Group B material (i.e., troublesome in fluidized dense-phase causing severe pipe vibrations). It was recommended to the power station that the baghouse operating sequence be changed in such a way to ensure the blow tanks were filled only during the cleaning cycle. This change was made easily via the control system and the plant has been operating successfully since this time. 2.2

Blow Tank Design

Some typical “conventional” blow tank designs are illustrated in Fig. 7. This type of feeder has been used successfully in industry to handle a wide range of products over relatively short distances (e.g., L = 100 to 200 m). However, the application of this type of blow tank to greater distances and/ or capacities (usually in conjunction with a second blow tank unit in series or parallel to ensure an essentially continuous mode of flow) resulted in a number of problems. • Incomplete and/or inefficient discharge of material from each blow tank mainly due to rat-holing or funnel-flow effects promoted by — low pressure drop across the bed of material (under steady-state conditions),

11-Oct-2001 JMR


(a) (a)

Conveying Material

Air..

Pipeline

Inlet ~

Vent :i$:

Discharge Valve

::8:

""'

Porous Membrane or Gas Distributor

\,.1 Aeration

Air+

(b) (b) Figure 7. “Conventional” blow tank feeders: (a) bottom-discharge arrangement; (b) top-discharge arrangement.

11-Oct-2001 JMR


735

— cohesive properties of the bulk solid, and/or — inappropriate method of blow tank air injection (e.g., single-point injection, excessive top-air promoting material compaction), • Flow instabilities or even pipeline blockage during start-up procedures due to surge effects caused by the initial high pressure drop across the blow tank (and hence high instantaneous flow rate of material), • Pipeline blockage during shutdown procedures due to surge effects caused by the sudden venting of high pressure air down the pipeline (and hence, acceleration of product towards the end of the pipeline). To avoid these problems, a number of significant developments have occurred in the area of blow tank design, such as: • Method of air injection (Kennedy, et al., 1987), in particular the fluidizing-discharge-cone (Wypych, 1995a). • Cone-dosing valve (Cürten, 1982) to control and meter the product into the pipeline. This is achieved by the movement of a double cone in the vicinity of the blow tank outlet. The cone moves continuously up and down inside the vessel but the stroke is adjusted by a proportional-integral (PI) controller based on a conveying line back-pressure measuring signal and set-point. The higher the conveying or operating pressure, the lower the cone set point inside the vessel thus restricting the flow of solids into the conveying pipeline (until a satisfactory conveying pressure is restored). Furthermore, the cone dosing system provides an additional and useful feature of mechanical agitation (and hence, flow assistance) of materials possessing cohesive and/or poor flow properties, such as manganese oxide (Wypych, 1989a). These developments have resulted in a much more reliable and efficient blow tank feeding system, especially for long-distance applications, as shown in Fig. 8. Such systems now have been used successfully in many installations throughout Australia to meet the increasingly demanding requirements of conveying capacity and distance.

11-Oct-2001 JMR

736 Fluidization, Solids Handling, and Processing Actuator Vent

Material Inlet

Actuator + ro--o

J.-

Vent

Top Air

Too Air

T. ~

-j

Cone Dosing Valve AeratIon Air

/-

...~

?r

..Fluldlslng -"--

/

Discharge

~

?, ,

Cone Conveying

Air .

Pipeline

Figure 8. Preferred tandem blow tank feeding system for long-distance pneumatic conveying.

Note, other options to control the discharge of material from the blow tank feeder include a rotary valve and an oscillating or modulating valve (Marcus, et al., 1980), as shown in Figs. 9 and 10, respectively. However, the former option requires special start-up and shutdown procedures (e.g., minimizing pressure drop across the valve) and is limited to low-temperature and low-abrasive applications. Also, the valve needs to be designed to handle high static pressures and will add to the capital and maintenance cost of the system. The latter option of using a modulating valve will increase maintenance costs, especially if the same valve is used for blow tank isolation. It should be realized also that either option shown in Fig. 9 or 10 will produce a non-symmetrical flow pattern inside the blow tank (i.e., due to preferential feeding at the blow tank outlet) and hence, promote the possibility of arching, rat holing and/or formation of dead regions. For these reasons, the combined fluidizing-discharge-cone and cone-dosing valve system shown in Fig. 8 is preferred.

11-Oct-2001 JMR


737

Figure 9. Rotary valve feed rate controller.

Figure 10. Modulating valve controller.

Case Study. As a part of the initial design of two 350 MW units for a coal-fired power station, it was required to: • collect all the ash from the electrostatic precipitators and economizers,

11-Oct-2001 JMR

738 Fluidization, Solids Handling, and Processing • transfer this material at ms ≈ 100 t h-1 over a total distance of L ≈ 1.5 km to a final storage silo, and • provide a fly ash resale facility outside the boundaries of the power station (i.e., to enable easy access for the contractors to pick up and deliver the ash to various cement plants). Tandem 7 m3 bottom-discharge cone-dosing blow tanks, similar to those shown in Fig. 8, were employed for this purpose and also to regulate the feed of material into the pipeline and alleviate the occurrence of imminent blockages (Wypych, 1995b). 2.3

Conveying Characteristics

When it is necessary to design or evaluate a proposed pneumatic conveying system, it is recommended strongly that the designer obtain as much information as possible on the actual material(s). Armed with the knowledge of steady-state conveying characteristics, it is quite a simple task to determine the minimum conveying velocity, optimal operating conditions for the product, and the pipeline diameter and compressor/blower rating to suit a given ms and L. Alternatively, conveying characteristics may be used to investigate operational problems that an existing plant may be experiencing (e.g., frequent blockages, reduced conveying rates). Problems of excessive product degradation and/or system wear also may be minimized by using such information to establish a safe minimum value of mf for a given ms. That is, conveying characteristics will determine whether an existing plant is operating at an optimal condition. If not, they will reveal what modifications would be necessary to achieve the desired result. Conveying characteristics also will provide useful information when an existing plant needs to be upgraded to achieve say, a higher conveying rate of solids. For example, it will be possible to determine whether the system and the material will be able to cope with the increased pressure and/or air flow requirements (i.e., whether the combination of pipe size and blower/compressor rating will be sufficient). The determination of steady-state conveying characteristics for a given product and test rig has been the subject of a number of earlier investigations such as Mason et al. (1980) and Mills et al. (1982). A standardized test procedure also has been developed and presented by

11-Oct-2001 JMR


739

Wypych and Arnold (1985b) and hence, only a brief description is presented here. The test procedure basically consists of three different types of experiments which are applied to the material until sufficient data have been collected for the determination of conveying characteristics. The steady-state parameters generated specifically for this purpose are • mf = Supplied mass flow rate of air (kg s-1), • ∆pt = Total pipeline air pressure drop (kPa), • ms

= Mass flow rate of solids or conveying rate (kg s-1 or t h-1).

Some typical examples of pneumatic conveying characteristics for three different fly ash samples conveyed on the same long-distance test rig are presented in Figs 11, 12 and 13. Some important information regarding these materials and results is summarized below. • Fly Ash A1 (Fig. 11): d v50 = 5 µm, ρ s = 2540 kg m-3, ρ bl = 670 kg m-3, Geldart (1973) and Dixon (1981) Group C material. • Fly Ash A2 (Fig. 12): d v50 = 75 µm, ρ s = 2500 kg m-3 and ρ bl = 1175 kg m-3, Geldart (1973) Group A material but Dixon (1981) Group B material. The latter was confirmed by the nature of the blockages obtained on the test rig (e.g., severe pipe vibrations), as well as the rapid observed deaeration rate of the material. • A company was intending to transport Fly Ash A2 through a long-distance conveying system that was designed to handle 45 t h-1 of Fly Ash A1. From Figs 11 and 12, it can be seen quickly that A2 would have required almost twice as much air than A1 and hence, a significant increase in running costs and possibly wear. • Fly Ash B (Fig. 13): d v50 = 12 µm, ρ s = 2215 kg m-3 and ρ bl = 955 kg m-3, Geldart (1973) and Dixon (1981) Group C material. It is interesting to see how much lower the pressures are for this material compared with Fly Ash A1. The above three conveying characteristics demonstrate the wide range of performances that can occur for fly ash and hence, the importance of employing such information in the design and optimization of longdistance systems.

11-Oct-2001 JMR


Figure 11. Test rig conveying characteristics of fly ash A1, L = 945 m, D = 69/81 mm and r = 1 m.

-..600 .

~ ~

g.

.

500

.. Q

.. ~

400

/' Blockage

~

..300 .5 "Gi Q. A: 200 -; .. 0

~

ms

8.4;-(t/h) ..

.--:..-=---

.....

..

-~6 ... 4

f

Vfi=9m1s,

2 ~

~ ..

..

1000.05 I

0.10 I

0.15 I

Air

Mass

0.20

Flow

Rate

0.25

0.30

(kg/s)

Figure 12. Test rig conveying characteristics of fly ash A2, L = 945 m, D = 69/81 mm and r = 1 m.

11-Oct-2001 JMR


741

Figure 13. Test rig conveying characteristics of fly ash B, L = 945 m, D = 69/81 mm and r = 1 m.

2.4

Pressure Drop Prediction

For long-distance and/or large-throughput applications, it is essential to predict accurately the total pipeline air pressure drop, ∆p t. For example, even with a “good” existing model predicting say, ∆p t = 500 kPa ±30%, the resulting uncertainty still would be too great (i.e., 350 to 650 kPa) and could lead to serious operating problems (e.g., inadequate capacity, pipeline blockage). Also, various major deficiencies exist with current models, such as the following: • Existing correlations usually are based on low values of air density, whereas ρ f can have a significant effect on ∆p t • Products that are suited to long-distance transportation usually have a wide range of particle size (e.g., fly ash, d = 1 to 300 µm), and it is difficult to represent adequately such materials by the single diameter required by most models • Most existing models are suited only to “pure” dilute-phase or dense-phase (high m*) applications, whereas the possible modes of flow over long distances occur between these two extremes (e.g., dune-flow, sliding beds, irregular slugging, etc.—usually at moderate m*).

11-Oct-2001 JMR

742 Fluidization, Solids Handling, and Processing For these reasons and other complex influences (e.g., large-diameter pipelines, particle-wall friction, particle shape, bends, etc.), it has been accepted that if high accuracy is needed, then some form of empiricism must be adopted. The preferred test-design procedure is listed below. • Firstly, as described previously, characterize the bulk solid to be conveyed by undertaking particle size, particle density, loose-poured bulk density, fluidization and deaeration tests. Then classify the material using both the Geldart (1973) and Dixon (1981) diagrams to establish possible feeding problems and dense-phase suitability, which provides an indication of long-distance performance. For example, based on experience, Dixon (1981) Group A (and some Group C) powders display good fluidized dense-phase (Wypych, 1995a), as well as longdistance performance (e.g., high m*). • Using a standardized test procedure (Wypych and Arnold, 1985b), determine test rig conveying characteristics, similar to those shown previously in Figs. 11–13. • Where the required conveying lengths and/or diameters cannot be tested, appropriate scale-up procedures (Pan and Wypych, 1992a) are employed. By monitoring the pressures along the straight sections of a test rig pipeline and also monitoring bend effects, determine exponents x1, …, x4 and y1, …, y4 in the following pressure drop equations. Eq. (7)

∆p s = ∆p sf + ∆p ss

where ∆p sf can be determined from Wypych and Pan (1991) and 2 ρ fm V fm

Eq. (8)

∆p ss = λ ss m ∗

Eq. (9)

λss = x1 m*x2 Frmx3 ρ fmx4

Eq. (10)

∆p b = ∆p bf + ∆p bs

2D

∆L s

where ∆p bf can be determined from Wypych and Pan (1991) and

11-Oct-2001 JMR


∗

743

ρfoV fo 2

Eq. (11)

∆ pbs = λbsm

Eq. (12)

λbs = y1 m*y2 Froy3 ρ foy4

2

Note that the exponents x1, …, x4 and y1, …, y4 are valid only for the test material and bend geometry, respectively. To demonstrate the scale-up accuracy of the above design equations, thirty-eight experiments were carried out (Pan and Wypych, 1992a) with a particular fly ash over a very wide range of conveying conditions (i.e., from dilute- to fluidized dense-phase) on the test rig Pipeline I shown in Fig. 14. In each experiment, it was believed that all the transducers along the pipeline (Tb-Te) were installed beyond any bend effects. Based on the data obtained from these experiments, the exponents in Eqs. (9) and (12) were determined by minimizing the sum of the squared errors of pressure at points Te2, Tc1, Tc2, Tc3 and Tc4, starting from point Te1. The determined values of exponent are listed in Table 1. Additional experiments then were carried out on Pipelines II and III, which are shown in Figs 15 and 16 respectively and detailed in Table 2.

13.0m

Loca .

.OC ~ I' ~ :

000

Transducer

Uon

: 6.7m

.I Q ~ O

I I I

~

t...000Tel

~~~~~

Te2 0

.

."' 0, ~ 7.6 m

',,- "

,

Tbl',

~

5.9 m

~ Tc4

~ Tc3

~ 73.5 m

0 Tc2

, "

~-, Tcl

Figure 14. Schematic layout of Pipeline I, L = 172 m, D = 69 mm, r = 1 m.

11-Oct-2001 JMR

744 Fluidization, Solids Handling, and Processing Table 1. Values of Exponent Based on Data Obtained from Pipeline I, L = 172 m, D = 69 mm and r = 1 m Exponent

Value of Exponent

x1

5.306

x2

-0.436

x3

-1.934

x4 y1

-0.117 0.052

y2

0.658

y3

0.673

y4

-1.495

---, 1

= 5

I1

9

: 16.3m I 1 1 : ~ ~,-~ " ~~

.I O ~ O ~

'0'

Transducer Location

0

0

, , ., , , , .J.' Tel '"',. ',. '. "

,

"

~7.6m

Tbl..

.'

0 ~

0 Tc4

0--~-~

... ~~1~~

.. ~~..

'--

'

~ 59.m

' , , , '.,'

., .. To3 ToZ.,...To1.

',' "

"

..,

T= -"

.,'

.'

, " ," , ' " ~ ,",.',

~59.0 m

'0' Tc3

73.5 m

0 Tc2

'0'.. Tcl

Figure 15. Schematic layout of Pipeline II, L = 554 m, D = 69 mm and r = 1 m.

11-Oct-2001 JMR


13.0m w--M

745

... ..... ... .... D = 69 mm 6.1 m Trl ‘*‘-““‘-----y f ----

----D=81mm

--59.0m

- - - - - 7

l%1’‘a ..-....-......o..-..-.-..........-*..............-..................~............*..-Tc2 Tel Tc4 Tc3 73.5m Figure 16. Schematiclayout of PipelineIII, L = 945 m, D = 69/81mm andr = 1m.

Table 2. Configuration of Test Rig Pipelines

Using the above exponents and pressure drop equations, the total pipeline air pressure drop and selected pressures were predicted for each Pipeline I, II and III, starting from the end of pipeline or points along the pipeline. Some of the predictions are presented in Fig. 17, from which it can be seen that the agreement between predicted and experimental values is quite good.


Experimental value (kpag) (a) Pipeline I (L=172 m).

Experimental

value (kpag)

(b) Pipeline II (L=554

m).

Experimental value (kpag) (c) Pipeline III (L=945 m).

Figure 17. Predicted pressure vs experimental value, based on data from pipeline I. (Note: Te1 → Tc4 means the predicted pressure at Tc4, starting from Te1).

11-Oct-2001 JMR


747

The above test-design and scale-up procedures have been applied to numerous long-distance systems (e.g., fly ash, pulverized coal) and have been found to provide good accuracy and reliability. Some examples of these results have been presented by Pan and Wypych (1992a) and Pan and Wypych (1994). The interested reader is referred to these papers for further details. The above procedures also have been used successfully to • Predict conveying characteristics for a wide range of pipeline configurations (Pan and Wypych, 1994) • Investigate the effect of bend number and radius on conveying performance (Wypych and Pan, 1993) • Compare the pressure drop caused by horizontal and vertical sections of pipe (Wypych and Pan, 1993) • Predict the pressure drop of large-throughput cement pneumatic conveying systems (Wypych, 1992) 2.5

Stepped-Diameter Pipelines It is recognized commonly that

Eq. (13)

Erosion ∝ Velocity n

where the power index, n, ranges from 2 for ductile materials to 6 for brittle materials (Marcus et al., 1990). Hence, one direct way of reducing pipeline wear (e.g., pipe, bends) is to limit the “natural” increase in velocity in the direction of flow (i.e., due to the expansion of the carrier gas). This can be achieved by increasing the bore of the conveying pipeline in the direction of flow. Other advantages of stepped-diameter pipelines include the minimization of pressure loss, air flow and hence, power consumption, which are particularly important when considering long-distance and/or large-throughput applications. By selecting accurate stepping pipe criteria and models to predict pressure drop (Wypych and Reed, 1990), it is possible to optimize the design of these pipelines and obtain more efficient transportation over longer conveying distances (e.g., up to L = 3 or 4 km). Some examples of long-distance stepped-diameter pipeline systems include 100 t h-1 of fly ash over 1.5 km (Wypych, 1995b) and 24 t h-1 of pulverized coal over 1.8 km (Wypych et al., 1990). Some of the different pipeline configurations considered for the latter are repeated in Table 3 below. Note:

11-Oct-2001 JMR

748 Fluidization, Solids Handling, and Processing • The relationship “Frmin = 6 = constant” was found to represent adequately the reliable transport limit of this material and hence, was employed as the pipe stepping criterion (i.e., to the optimize Ln and minimize mf), • An accurate λs correlation based on experimental data was employed to predict the values of ∆p t. From Table 3 it can be seen that by optimizing the configuration of pipeline, it is possible to reduce pressure loss, air flow, transport velocity, and hence, pipe/bend wear. Depending on hardware requirements and reliability, which would to some extent govern the maximum operating pressure of the system (e.g., say, 400 or 500 kPag), Pipeline Nos. 5 or 6 could be selected for this long-distance application. However, if diverter valves are required at the end of the pipeline, it may be more convenient to select Pipeline No 5 (i.e., D1 = 154 mm instead of 203 mm). Table 3. Pipeline Configurations and Predicted Operating Conditions for 24 t h-1 of Pulverised Coal over 1.8 km (Wypych, et al., 1990) Pipe D n No (mm)

Ln

mf

m*

(m) (kg/s) (-)

∆pt

Fri

V fi

V fe

P

% Diff

(kPa) (-) (m/s) (m/s) (kWh/t/km) wrt No 6

1

203

1800 1.22 5.5

272

6

8.5

31.3

6.1

97%

2

154

1800 0.87 7.7

430

6

7.4

38.8

5.5

77%

3

127

1800

6

6.7

46.1

5.2

68%

203

853

6

8.5

18

154

947

6

7.4

14.7

3.9

26%

154

1117

6

7.4

25

127

683

6

6.7

10.8

3.6

16%

203

436

6

8.5

13

154

733

6

7.4

14.7

3.1

0%

127

631

6

6.7

10.8

4 5

6

2.6

0.7

9.6

595

0.7

9.6

327

0.56

12

453

0.51 13.2 401

Valves

Blow tank technology was introduced extensively after World War II to not only improve plant efficiency but also to extend the future potential of pneumatic conveying (e.g., long-distance and/or large-throughput

11-Oct-2001 JMR


749

applications). However, the resulting higher operating pressures, together with other factors such as product abrasion and/or high temperatures, resulted in the premature failure of crucial valves (e.g., blow tank inlet, blow tank discharge). A wide range of different “off-the-shelf” valves for gas, hydraulic and/or slurry service then were tested (e.g., ball, butterfly, plug and air-on-sleeve pinch valves). Unfortunately, most of these were found unsuitable, achieving service lives from only a few hours to a few days. Consequently, more elaborate designs had to be tested and/or developed for this application of pneumatic conveying. The following two valves have been found useful in particular areas. Note that in each case it is imperative to ensure fast actuation time (e.g., 1 second for a 100 mm nominal bore NB valve and a few seconds for a 300 mm NB valve, if possible). Large-bore solenoid valves and quick-exhaust valves usually are required for this purpose. Also, note that each one of the following valves provides a full cross-sectional area of flow in the open position and a 100% seal in the closed position. One of the major problems of air-on-sleeve pinch valves is that they do not provide these important features for pneumatic conveying applications (e.g., a small hole in a closed sleeve quickly erodes due to the subsequent high velocities of air and solids). The high-pressure pinch valve shown in Fig. 18 has been applied successfully to several abrasive pneumatic conveying applications, mainly for blow tank discharge. For example, this pinch valve has been quite successful in handling abrasive and coarse materials, such as crushed bauxite, zircon and sub 20 mm crushed brick (Wypych, 1995b), as well as high throughputs, such as 100 t h-1 of cement (Wypych, 1995b). Also, by replacing existing chrome-plated butterfly valves with high-pressure pinch valves in cement plants, it has been possible to extend the service life of blow tank discharge valves from ≈2 weeks to over 8 months (Timms, 1992). However, it is important to select a valve which has the actuating rods “tagged” or connected to the sleeve. This will ensure a full-bore flow area when the valve is opened. Note that this type of valve is not suitable for high-temperature applications (e.g., ≥200 °C), where the sleeve will work-harden and fail prematurely. The rotating-disc valve shown in Fig. 19 is gaining popularity for blow tank inlet and discharge applications. As the disc is spring-loaded and “selflapping,” the valve automatically adjusts for wear. Also, as no packing, seals or o-rings are used (i.e., the disc and seating surfaces are metallic),

11-Oct-2001 JMR


Actuator

/

Flanged Sl eeve Throttle

Bar

Valve

Body

Throttle 1 Valve

Bar-r-

~

Flange

'-

""

I

'.1i;::9/ Valve

Open

Closed

Figure 18. High-pressure pinch valve (Larox Pty Ltd).

I NO.I ~ c

3 4

NAME Body

Plate.

DESCRIPTION I

wrench Post Lever

Carbon

Steel

Carbon

Steel

A515

17-4 PH Stainless Carbon Bearings

Seml-

6

Disc

Cast Ste111le .6

I

Disc

8

I

Seat

g

I

Spacer

Spring

I

17-7

70

Sleel

Steel A515 Gr. 70

5

7

Or.

Steel

PH

Slolnloss

Slool

I Cast Stelllle #6 Ring

I

Carbon

I 10 I Gaskets

I

Corrugated

Steel

A106

r--11'-P8cklng

I J. Crane 287-1

304

Gr.

Stainless

C Steel

Nllronlc 60 17-7 PH SS Bellevllle Washers

20

FIO.00..T".'.

Figure 19. Rotating-disc valve (Everlasting Valve Co.).

11-Oct-2001 JMR

6 314-'-6 112 'FIG.10).IT1..'"


751

high temperatures and pressures are possible (e.g., up to 900°C and 69000 kPag). The rotating-disc valve has been used successfully (Wypych, 1995b) as a discharge valve for 7 m3 blow tanks transporting hot and abrasive ash from a power station precipitator at a rate of 100 t h-1 over a total distance of 1.5 km. These discharge valves have operated continuously on a 24-hour basis and achieved a service life of at least 12 months (i.e., without any form of maintenance). However, to ensure that the valve body does not fill up with ash, purge-air has been used during actuation of the valve. 2.7

Pipeline Unblocking Techniques

Following numerous blockages on the test facilities at the University of Wollongong (attempting to determine the minimum conveying velocity of various materials over long distances), a pipeline unblocking system was developed and installed at the end of the pipeline. Refer to Fig. 20 for a typical arrangement. The system incorporates only one component in the pipeline, allowing the pipeline to be back-pressurized from the silo to the blockage with air at a slightly higher pressure than the conveying pressure. The back-pressure then is released in a controlled manner such that the blockage is drawn through the conveying line. These systems have been incorporated successfully in the control circuits of plants handling difficult-to-convey materials and also have been used on conveying pipelines up to 1.5 km in length (Wypych, 1995b).

Back.Pressure Valve"

Figure 20. Schematic layout of pipeline unblocking system.

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General Considerations

The following comments and recommendations are based on the research findings presented and cited in this paper, as well as the consulting experiences of the author. For the general purpose of minimizing air flow, transport velocity, wear and power, the fluidized dense-phase mode of flow is preferred for long-distance applications. Efficient blow tank feeders, rotary-screw compressors, refrigerated dryers and stepped-diameter pipelines also are recommended. For products that are not suited to fluidized dense-phase, the possible modes of flow include dilute-phase (suspension flow) or bypass conveying (Wypych, 1995a). It is believed that the air velocities in a large-diameter dilute-phase system can be 50 to 100% higher than an equivalent well-designed densephase system. Hence, much greater wear problems are expected in the dilute-phase system, although significant advances have been made in the technology of wear-resistant materials and bends (Wypych and Arnold, 1993). Other features involved with dilute-phase transport systems include: • Limited operating pressures (e.g., 80 to 100 kPag for Rootstype blowers) • Large-diameter pipelines • Greater conveying velocities and system erosion (i.e., due to the larger pipe diameters) • Larger dust collectors to cope with the greater volumes of conveying air Note, the velocities required in large-diameter dilute-phase systems must be high to ensure suspension flow, avoid product deposition, and facilitate good clean-out. This comment is based on personal experiences involved with large-diameter pipeline systems used to unload ships (e.g., 350 mm NB pipelines 25% full of deposited cement) and provide fuel to boilers (e.g., deposition of pulverized coal even at high velocities such as 30 m s-1). Note, blow tank feeders (i.e., instead of rotary valve feeders) can be used effectively to assist in the cleaning/purging of pipelines (e.g., cleanblow cycle). However, with a Roots-type prime mover and a largediameter (large-volume) pipeline, the effectiveness of the purge or cleanblow cycle will be reduced significantly (i.e., due to the pressure limit of 100 kPag).

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753

In contrast, a well-designed dense-phase system will operate at significantly less velocities and wear, as well as provide more efficient cleaning/purging operations (i.e., due to having smaller pipe diameters and greater pressures, if needed). One other interesting issue is that a blow tank dense-phase system is inherently more “self-cleaning” than a dilute-phase system (i.e., due to the natural variations in pressure and increased material concentration/turbulence). Quite often dense-phase systems have been installed in Australia to overcome buildup problems displayed by dilutephase systems, where the air can flow easily over deposited material. When evaluating a material for the purpose of establishing densephase and long-distance suitability, it is important to undertake all the necessary tests (e.g., particle sizing, particle and bulk densities, fluidization and deaeration). Also, if possible, it is useful to compare such results with those obtained on previously conveyed similar materials (e.g., fly ash). However, it should be noted that such an evaluation only is a qualitative one and it is not possible to predict say, minimum air flows or pipeline pressure drop based on such data (i.e., pilot-scale tests normally are required to confirm minimum velocities, friction factors, etc., especially over long distances and for large-diameter pipes). Many designers and researchers in the past have placed a great deal of emphasis on the importance of solids loading or the solids-to-air mass flow rate ratio, m*. However, it should be realized that this parameter is dependent on particle and loose-poured bulk densities; conveying distance and pressure available; frictional properties and minimum transport conditions of material (which will affect the maximum values of m*). Hence, it is difficult and misleading to apply “general” values of m* to different systems and materials. The objective always should be to achieve a welldesigned long-distance transport system operating at optimal conditions— m* only should be considered as a consequence of achieving this result.

3.0

PIPE BRANCHING

The branching of pipes/ducts in pneumatic conveying has the following two main applications in industry. • Extraction of dust via a network of branched ducts, where the particulates are transported under vacuum conditions. Some common examples include the control of dust in materials

11-Oct-2001 JMR

754 Fluidization, Solids Handling, and Processing handling operations, including conveyor transfers, screening, filling bins, ship loading, etc. • The simultaneous splitting and distribution of solids-gas mixtures for applications requiring multipoint injection, where the mixtures are transported usually under positive-pressure conditions. Some common examples include tuyere injection for blast furnaces, large burner nozzles for pulverized coalfired boilers, small coal-fired plasma torches providing startup and support energy for boilers, injection of pulverized fuel into calciners, etc. Some of the important issues that should be considered when designing, improving or operating any such pipe branching applications are described in the following sections. 3.1

Dust Extraction

Most practical dust extraction systems involve multiple hoods linked to a duct network serving a central gas cleaning unit and prime mover. This is very similar to an air conditioning system where the air is supplied from an air handling unit through the duct network to diffusers serving the occupied spaces. In fact, most duct sizing methods are based on air conditioning principles and/or techniques, such as the Total Pressure Method (ASHRAE, 1985), the Velocity Pressure Method (ACGIH, 1992) and the Constant Pressure Gradient Sizing Method (DASCG and AIRAH, 1987). Also, some “air-conditioning” computer programs, such as DONKEY (DASCG and AIRAH, 1988), may be employed to assist in the design of the ductwork (i.e., based on one or two of these sizing methods). The objective of the design/analysis process is to ensure (as far as possible) that the correct quantity of air flows through each hood. As different air quantities, hood types/sizes and/or duct lengths usually are involved, this is not always a simple task. The correct air flows may be achieved by performing: • A “rough” duct design and then making adjustments during commissioning using blast gates or dampers to control air flow • A more involved design where the system is inherently balanced. The latter method is preferred for the following reasons:

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755

• Dampers may increase maintenance costs • Dampers may be tampered with during the life of the plant • Damper control may be prohibited when handling dangerous or explosive materials • The “rough” duct design method may result in premature dust deposition or excessive system erosion (due to incorrect duct sizes and velocities) Hence, even if the “rough” duct design technique is selected, some form of balancing still should be performed so that the final duct sizes are close to optimal (in terms of pressure and velocity). The correct application of either sizing technique will result in a duct network that works well on air. It is equally important to ensure that each air flow is adequate to transfer all the particulates from the hood/enclosure to the collection and/or cleaning device. Unfortunately, many dust control systems have been designed and/or are being operated with little or no regard for what actually has to go through the hood-duct network. This can result in: • Particulate deposition and buildup inside the duct(s) • Eventual choking of duct branch(es) • Inefficient performance of the overall dust control system. Case Study. Such a situation was found to occur in the duct network shown in Fig. 21 and installed to extract iron oxide dust at various points along a cold strip processing line. The stated problems were insufficient suction at the hoods, buildup of contaminant in the hoods and along the processing line (causing cleanup problems due to eventual mixing with hydraulic fluid, lubricant, water, etc.). Analysis of the system found the following: • Branch II-V contained more bends and was over three times longer than branch I-V • Branches I-V and II-V were sized to provide an approximate 35%–65% split-up in air flow, respectively (i.e., using pressurebalancing techniques) • The system would have worked initially on air, however, the lower section of branch II-V was completely blocked with contaminant

11-Oct-2001 JMR

756 Fluidization, Solids Handling, and Processing • Velocities in excess of 60 m s-1 were occurring in branch I-V causing excessive pressure loss and wear. Note the potential seriousness of wear in ductwork (especially bends/elbows) has been emphasized previously by Eq. (13). For example, assuming n = 3, which without other data is used to represent most industrial situations, a velocity increase of only 20% will result in a 73% increase in wear. Hence, it is important to ensure that the transport velocities are not only adequate to avoid deposition problems but also minimized for wear reasons. The solution to the above problem is to re-size the entire network, in particular branches I-V and II-V, with the aim of ensuring a suitable minimum transport velocity along each section of duct.

Figure 21. Duct network for extraction of iron oxide dust.

Minimum Duct Velocity. Once the types and locations of hoods, enclosures and booths have been established and the exhaust rates determined, it is necessary to design the ductwork to transfer the contaminated gas to the collection or cleaning device. An essential parameter required for this purpose is the minimum duct or conveying velocity, Vfmin , which is required to:

11-Oct-2001 JMR


757

• Avoid the gradual buildup of a particular contaminant with time • Know whether any toxic or bacteria-prone particles are being left inside the ductwork (e.g., for servicing or maintenance purposes) • Establish whether any appreciable buildup will affect dust extraction efficiency Unfortunately, the current selection of Vfmin is based on a general list, such as that given in Table 4, where: • The list of contaminants is very general and open to interpretation • The suggested velocities can vary considerably between say, 15 and 20 m s-1 (even for the same application area) • No consideration is given to the influence of particle properties and/or duct diameter Table 4 Minimum Conveying Velocities (ACGIH, 1992) Vfmin Nature of Contaminant Vapours, gases, smoke Fumes Very light fine dust

Industrial Examples

-1

-1

(ft min )

(m s )

All vapours, gases and smoke

1000-2000

5-10

Welding Cotton lint, wood flour Fine rubber dust, bakelite moulding powder dust, cotton dust, light shavings, soap dust, leather shavings

2000-2500 2500-3000

10-12 12-15

3000-4000

15-20

Average industrial dust

Grinding dust, dry buffing lint, coffee beans, granite dust, silica flour, general materials 3500-4000 handling, brick cutting, clay dust, foundry, limestone dust, asbestos dust

18-20

Heavy dusts

Heavy/wet sawdust, metal turnings, foundry tumbling barrels and shake-out, sand blast dust, wood blocks, brass turnings, cast iron boring dust, lead dust

4000-4500

20-23

Heavy or moist

Lead dust with small chips, moist cement dust, asbestos chunks from machines, sticky buffing lint

> 4500

> 23

Dry dusts and powders

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758 Fluidization, Solids Handling, and Processing For example, for the iron oxide dust considered in the previous case study, Table 2 suggested Vfmin = 18 to 20 m s-1 (i.e., assuming an “average industrial dust”). On analysis of the sample, it was found d p50 ≈ 80 µm, which appeared to support this classification. However, upon further examination of the actual distribution of size, a significant proportion of the material was found > 1000 µm (e.g., large flakes). A minimum conveying velocity of at least Vfmin ≈ 25 m s-1 was estimated for this “dust.” This explains why the iron oxide material built up and eventually blocked branch II-IV, which was sized/balanced mainly for air distribution purposes and produced transport velocities < Vfmin. As a result of some recent investigations into the deposition (saltation) of small particles in large diameter ducts (Cable, 1994; Miletich, 1994) and analogous work in pneumatic conveying (Wypych and Reed, 1990; Zenz, 1964; Cabrejos and Klinzing, 1994), some other interesting characteristics of Vfmin are listed below. • Vfmin appears to increase with duct diameter, D. This trend has been recognized in many “traditional” pneumatic conveyors (Wypych and Reed, 1990; Zenz, 1964; Cabrejos and Klinzing, 1994), but also appears to be relevant for dust extraction (Cable, 1994; Miletich, 1994), which can be considered as a form of low-concentration pneumatic conveying (e.g., Vf = 10 m s-1 may be suitable for a particular contaminant in a duct size of D = 100 mm but may cause deposition in D = 300 mm). Further evidence of this can be seen in the work of Zenz, 1964, who found that single-particle saltation, which is analogous to the deposition of low-concentration particulate suspensions in dust extraction, increases with duct diameter according to the following relationship. Eq. (14)

Vfso ∝ Dk where k = 0.4 to 0.6. However, based on experience, Eq. (14) appears to provide conservative values of Vfso, especially in large-diameter ducts. This has been confirmed by Cable (1994) and Miletich (1994) who both employed 200, 300 and 480 mm diameter steel ducts, and Cabrejos and Klinzing (1994) who used a 50 mm diameter copper tube.

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759

• Vfmin appears initially to decrease with decreasing particle size and then increase for very fine particles (due to cohesive and interparticle forces). This trend is predicted (Wypych, 1993) by three different saltation velocity models and also can be found in many industrial applications. For example, in an article describing the installation of a dust extraction system for very fine clay dust in a brick manufacturing plant (Anon, 1995), it was stated that a high velocity of 24 m s-1 was needed to maintain dust suspension—this velocity is approximately equal to the 25 m s -1 estimated for the (coarser and heavier) iron oxide dust described above. Note, this example also provides further support to the inadequacy of Table 4, which nominates a velocity of 18–20 m s-1 for “clay dust.” It should be noted that these unusually high velocities may be indicative of what would be needed to ensure “clean” ductwork (i.e., without any form of deposition), whereas in practice a small amount of deposited material is considered normal. Nevertheless, it is essential to ensure that such a buildup does not increase gradually with time. Also, in some applications, especially those handling dangerous and/or bacteria prone contaminants, it may be necessary to maintain “clean” ductwork (i.e., for safety, hygienic and/or maintenance reasons). • The velocity needed to re-entrain deposited particulates (Vfup) may need to be determined for particular applications. The velocity Vfup is being found (Zenz, 1964; Cabrejos and Klinzing, 1994) to be much greater than Vfmin (i.e., the saltation or deposition velocity) and again increase with duct diameter (Cabrejos and Klinzing, 1994). Also, similar to Vfmin, Vfup appears to increase with deceasing particle size, where very fine contaminants are more difficult to re-entrain (i.e., due to interparticle forces). The further development of accurate models to predict the above parameters (including the effects of particle properties and duct diameter) is being pursued currently, and considerable effort still is required before such models can be applied solely in design practice (i.e., without the need for experience or comparative data).

11-Oct-2001 JMR

760 Fluidization, Solids Handling, and Processing At this stage, operators of dust control systems should at least be aware of these various issues and their relative importance, which will vary from problem to problem. Hence, it is essential that the engineer or consultant responsible adopts a systematic approach to the solution of the problem in which all the relevant issues are considered and dealt with in the appropriate order. If accurate design/analysis data are necessary (e.g., Vfmin, Vfup), it is possible to obtain the information needed by undertaking the appropriate bench-type tests and large-scale test-design procedures, similar to the approaches employed for pneumatic conveying (Pan and Wypych, 1992a; Arnold et al., 1994). 3.2

Flow Splitting

The layout of a typical flow-splitting system is depicted in Fig. 22. For an even split-up and distribution of material, it is essential that each pipe branch offers the same resistance to flow. This can be achieved easily by ensuring that each pipe branch has the same length, type and location of bends, as depicted in Fig. 23. This preferred approach is more feasible when small diameters of pipeline are involved. However, the installation of identical pipe branches is not always possible. In such situations, the traditional approach has been to install additional pipes/bends in the lower resistance branches, such as branch Nos. 1 and 4 shown in Fig. 22. The problem with this approach is that bend pressure drop depends not only on conventional parameters, such as mf, ms, material properties, wall material, temperature, bend geometry, etc., but also on bend location (Wypych and Pan, 1993; Pan and Wypych, 1992b). Hence, simply ensuring each pipe branch comprises the same length of pipe and number of bends still can cause inaccuracies in flow splitting efficiency. By employing accurate test-design procedures (Pan and Wypych, 1992a), it is possible to model and design each pipe branch separately so that the system ultimately is well balanced. However, such a system may not be reliable over time due to uneven wear in the pipes/bends, changes in material property and/or on-site conditions. Splitters. The following common splitters have been introduced to improve the efficiency of flow splitting.

11-Oct-2001 JMR


No2

761

No3

Solids-Gas Flow

Figure 22. Traditional method of pipe branch layout for injection systems.

Solids-Gas Flow

Figure 23. Preferred method of pipe branch layout for injection systems.

11-Oct-2001 JMR

762 Fluidization, Solids Handling, and Processing • Conventional riffle box, as shown in Fig. 24, used normally for 2-way splitting in coal-fired furnace and boiler applications. The riffles plates will be subjected to wear and over time could affect flow splitting efficiency. • Cone splitter, as shown in Fig. 25, used in general injection applications for up to 8-way splitting and claimed (Hilbert, 1982) to achieve ±10% accuracy in splitting. It should be noted that such figures depend more on the design of the pipe branches downstream of the splitter, rather than the splitter itself. • Rotary splitter (Selves and Barnes, 1993), which can be used to provide up to 36-way splitting. Due to its intermittent operation, the pulsing flow in the branches downstream of the splitter would not be suitable for applications requiring a smooth and regular injection of material. The above devices are flow intrusive, subjected to wear and the splitter itself cannot control changing downstream conditions. A potentially more direct and efficient approach is to monitor the change in flow conditions downstream of a splitter (Barnes and Murnane, 1995) and employ active splitters (Selves et al., 1995) to control the split ratio of air and hence, material. Some of the active splitters being investigated and developed by Selves et al., 1995 include: • Modified riffle box, as shown in Fig. 26, which uses air injection to control the split ratio. • Induced swirl, as shown in Fig. 27, which imparts to the solids-gas flow a swirling action and also controls the rate and direction of rotation of swirl via tangential nozzles. However, the residual swirl that would occur in the downstream pipe branches may cause problems if several swirl inducers follow one another (Selves et al., 1995). This problem could be eliminated by introducing the dropout box splitter shown in Fig. 28.

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763

Figure 24. Conventional riffle box. (Selves et al., 1995.)

Figure 25. Cone splitter. (Hilbert, 1982.)

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Figure 26. Modified riffle box (Selves et al., 1995).

Figure 27. Control of split ratio using induced swirl and a Y splitter. (Selves et al., 1995.)

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765

Figure 28. Improved dropout box splitter (Selves et al., 1995).

General Considerations. Some other important considerations that should be made when designing a flow splitting system are listed below. • The splitter should provide a symmetrical split in all planes and preferably should be installed in the vertical plane (if possible). • The solids-gas flow upstream of the splitter should be uniform and regular. • Sufficient upstream pipe should be used to eliminate any flow separation effects caused by in-line components, such as bends and diverters (e.g., “roping” in dilute-phase). • Special care should be taken when selecting the mode of solids-gas flow. For example, flow separation and roping could occur even in very dilute-phase conveying systems (e.g., m* ≤ 1 for coal-fired boilers). Fluidized dense-phase also is possible for some systems and can offer many

11-Oct-2001 JMR

766 Fluidization, Solids Handling, and Processing advantages (e.g., reduced air flow, velocity, wear, power). However, this mode of flow can be more irregular than dilutephase (e.g., increased pressure fluctuations) and hence, produce flow splitting problems, even with active splitters. The objective here should be to achieve the minimum air flow that is needed to ensure smooth and consistent flow under all operating conditions. 3.3

Pressure Loss

For dust extraction systems, the concentration of solids usually is quite low. For this reason, the methods employed to calculate pressure loss are based on air-only conditions. Comprehensive information is available (ASHRAE, 1985; ACGIH, 1992) to assist the designer in estimating the pressure loss caused by pipe branches, ducts, elbows, etc. In contrast, the amount of material being conveyed inside each pipe branch of a flow splitting application is very high and hence, design cannot be based on air-only analyses alone. For example, Low et al., 1987 have proposed the following empirical relationship to determine the head loss of a pipe branch. Eq. (15)

K = Kf (1 + C mfa)

where K is the branch head loss for solids-gas flow, Kf is the branch head loss for air-only conditions and can be calculated theoretically from (Low et al., 1987), C = 0.22 and a = 1.27 appear to represent satisfactorily both 90°and 45°-branches with different branch diameters and products (Low et al., 1987). Further examples of empirically based pressure loss equations for Y-splitters of various angles and subjected to plastic pellets under different solids loadings have been presented and cited by Marcus et al., 1990. Good flow splitting design is dependent on the accurate prediction of the pressure drop caused by the various bends, branches and straight sections of pipe. This can be achieved by employing the above branch model(s), proven for the particular material and application, coupled with the accurate “pipeline” test-design procedure described in Sec. 2.4 of this chapter.

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767

NOTATIONS a C d d 50 dp d p50 d pm d pwm d sv d svm dv d v50 d vm d vwm D Fr g k K Kf L Lh Lv mf ms m* n NB P r t Vf Vfso Vfup V mb

Power index Constant Particle diameter, m Median particle diameter, m Arithmetic mean of adjacent sieve sizes, m Value of d 50 based on a sieve size distribution, m Mean particle size from a standard sieve analysis, Eq (1), m Weighted mean diameter based on a sieve analysis, Eq (2), m Diameter of a sphere with the same surface area to volume ratio as the particle, m Mean surface volume diameter, Eq (3), m Diameter of a sphere with the same volume as the particle, m Value of d 50 based on a volume diameter distribution, m Mean equivalent volume diameter, Eq (4), m Volume weighted mean diameter, Eq (5), m Internal diameter of pipe, m Froude No, Fr = Vf (gD)-0.5 Acceleration due to gravity, m s-2 Power index Pipe branch head loss for solids-gas flow Pipe branch head loss for air-only conditions Total effective length of pipe or section of pipeline, m Total effective length of horizontal pipe, m Total effective length of vertical pipe, m Air mass flow rate, kg s-1 Solids mass flow rate, kg s-1 Solids to air mass flow rate ratio, m* = ms mf-1 Power index Nominal bore Specific power, W h t-1 m-1 Centreline bend radius, m Time, s Superficial air velocity, m s-1 Single-particle saltation velocity, m s-1 Re-entrainment velocity, m s-1 Minimum bubbling velocity, m s-1

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768 Fluidization, Solids Handling, and Processing Vmf Minimum fluidization velocity, m s-1 x1, …, x4 Exponents X Factor (Mainwaring and Reed, 1987) y1, …, y4 Exponents ∆h Difference in height, m ∆Ls Length of straight section of pipe, m ∆M Mass percent of material contained in a given size range, % ∆M i Value of ∆M for size range i, % ∆p Pressure drop, Pa ∆p b Pressure drop caused by bend, Pa ∆p s Pressure drop caused by straight section of pipe, Pa ∆p t Total pipeline air pressure drop, Pa κd Deaeration factor (Mainwaring and Reed, 1987) κp Permeability factor (Mainwaring and Reed, 1987) λb Particle-wall friction factor in bend λs Particle-wall friction factor in straight pipe ρ bl Loose-poured bulk density, kg m-3 ρf Air density, kg m-3 ρs Particle density, kg m-3 ψ Particle sphericity Subscripts i Initial value (at beginning of pipe) e Final or exit value (at end of pipe) f Fluid (gas) m Mean value (based on average air density) min Minimum value n Value relating to pipe section n (starting from end of pipeline) o Value relating to bend outlet s Solids

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769

REFERENCES ACGIH, Industrial Ventilation: A Manual of Recommended Practice, 21st Ed, American Conf. of Governmental Industrial Hygienists, Inc., Cincinnati, USA (1992) Allen, T., Particle Size Measurement, Chapman and Hall Ltd, London, 2nd Ed (1975) Anon, “Dynamics of Gas/Solids Systems,” Pneumatic Handling of Bulk Mat., Part A: Fundamentals, Bulk Solids Handling Unit, Thames Polytechnic, London (1983) Anon, “Dust Suspended from Workplace,” PACE, Process and Control Eng., p. 56 (1995) Arnold, P. C., Wypych, P. W., and Reed, A. R., “Advances in the Design of Pneumatic Transport Systems,” Powder Handling & Processing, 6(1):9–21 (1994) ASHRAE, ASHRAE Fundamentals Handbook , American Society of Heating, Refrigerating and Air Conditioning Engineers, Inc, Atlanta (1985) Barnes, R. N., and Murnane, S. N., “The Sensing of Unbalanced Pulverised Coal Feed Rates at the Exit of Riffle Boxes in Coal-Fired Power Station Fuel Distribution Systems,” 5th Int. Conf. on Bulk Mat. Storage, Handling and Transportation, Newcastle, IEAust, Proc., 1:273-281 (1995) Cable, P. M., Design Requirements of Ind. Dust and Fume Extraction Systems, BE Thesis, Dept. of Mech. Eng., Univ. of Wollongong (1994) Cabrejos, F. J., and Klinzing, G. E., “Minimum Conveying Velocity in Horizontal Pneumatic Transport and the Pickup and Saltation Mechanisms of Solid Particles,” Bulk Solids Handling, 14(3):541–550 (1994) Cürten, H. J., “Concepts and Results of Pneumatic Conveying of Support Materials in German Underground Mining,” Pneumatech 1, Int. Conf. on Pneumatic Conveying Technol., Stratford-Upon-Avon, UK (1982) DASCG, and AIRAH, Mechanical Engineering Services Design Aids No. DA3, Air Conditioning Duct Design Manual, (F. Wickham, ed.), Dept Administrative Services Construction Group and AIRAH, Aust Govt Publishing Service, Canberra, Aust (1987) DASCG, and AIRAH, Mech. Eng. Services Design Aids No DA4, User Guide for the Computer Program DONKEY, (F. Wickham, ed.), Dept Administrative Services Construction Group and AIRAH, Aust Govt Publishing Service, Canberra, Aust (1988) Dixon, G., “How do Different Powders Behave?,” Bulk-Storage Movement Control, 5(5):81–88 (1979)

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770 Fluidization, Solids Handling, and Processing Dixon, G., “Pneumatic Conveying,” Plastics Pneumatic Conveying and Bulk Storage, Applied Science Publishers, London (1981) Geldart, D., “Types of Gas Fluidization,” Powder Technol., 7:285–292 (1973) Geldart, D., and Abrahamsen, A. R., (“Fluidization of Fine Porous Powders, Recent Advances in Fluidization and Fluid-Particle Systems,” AIChE Symp. Series, AIChE, 77(205):160–165, New York (1981) Geldart, D., Harnby, N,, and Wong, A. C., “Fluidization of Cohesive Powders,” Powder Technol., 37:25–37 (1984) Hilbert, J. D., “Multiple In-Line Splitting of Pneumatic Conveying Pipelines,” J. Pipelines, 3:161–172 (1982) Jones, M. G., and Mills, D., “Low-Velocity Pneumatic Conveying: Product Characteristics,” Interbulk 89, Seminar on Pneumatic Conveying: Potentials and Capabilities, Birmingham, Organised by Glasgow College, Glasgow, Scotland and Thames Polytechnic, London, England (1989) Kennedy, O. C., Wypych, P. W., and Arnold, P. C., “The Effect of Blow Tank Air Injection on Pneumatic Conveying Performance,” Pneumatech 3, Int. Conf. on Pneumatic Conveying Technol., Jersey, Channel Islands, UK (1987) Lohrmann, P. C., and Marcus, R. D., “The Performance of a Bottom-Discharge Blow Vessel Pneumatically Conveying Three Group A Materials,” Bulk Solids Handling, 4(2):409–412 (1984) Low, H. T., Winoto, S. H., and Kar, S., “Pressure Losses at the Branches of a Pneumatic Conveying System,” Bulk Solids Handling, 7(6):865–867 (1987) Mainwaring, N. J., and Reed, A. R., “Permeability and Air Retention Characteristics of Bulk Solid Materials in relation to Modes of Dense-Phase Pneumatic Conveying,” Bulk Solids Handling, 7(3):415–425 (1987) Marcus, R. D., Leung, L. S., Klinzing, G. E., and Rizk, F., Pneumatic Conveying of Solids, Chapman and Hall Ltd, London (1990) Mason, J. S., Mills, D., Reed, A. R., and Woodcock, C. R., “The Use of Product Conveying Characteristics in the Design of Pneumatic Conveying Systems,” Powder Europa. 80, Seminar D, pp. 58–80 (1980) Miletich, D., Atomisation and Pneumatic Conveying of Coarse Aluminium Powder, BE Thesis, Dept of Mech Eng, Univ. of Wollongong (1994) Mills, D., Mason, J. S., and Stacey, R. B., “A Design Study for the Pneumatic Conveying of a Fine Particulate Material,” Solidex 82, pp. C1–C75 Harrogate, UK (1982) Molerus, O., “Interpretation of Geldart’s Type A, B, C and D Powders by Taking into Account Interparticle Cohesion Forces,” Powder Technol., 33:81–87 (1982) Pan, R., and Wypych, P. W., “Scale-Up Procedures for Pneumatic Conveying Design, Powder Handling & Processing, 4(2):167–172 (1992a)

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Pan, R., and Wypych, P. W., “Bend Pressure Drop in Pneumatic Conveying of Fly Ash,” Powder & Bulk Solids Conf., Reed Exhibition Companies, USA, Proc, pp. 349–360, Rosemont, Illinois, USA (1992b) Pan, R., and Wypych, P. W., “Design of Economic Pneumatic Conveying Systems,” MECH’94, Int. Mech. Eng. Congress & Exhibition, IEAust, Proc, 2:137– 141, Perth, WA (1994) Selves, T. P., and Barnes, R. N., “A Review of In-Line Splitting Techniques used in Pneumatic Conveying,” 4th Int. Conf. on Bulk Materials Storage, Handling and Transportation, 2:353–358, IEAust. Proc, Wollongong (1993) Selves, T. P., Barnes, R. N., and Reed, A. R., “The Use of Flow Diverting Air Injection to Actively Control the Split Ratio of Pneumatically Conveyed Particulate Materials,” 5th Int. Conf. on Bulk Materials Storage, Handling and Transportation, IEAust. Proc, 1:263–271, Newcastle (1995) Timms, G., “Pneumatic Conveying - a New Application for Larox Pinch Valves,” Larox News, 14:8–11 (1992) Wypych, P. W., Pneumatic Conveying of Manganese Oxide, ITC Bulk Materials Handling Report for BHP Engineering, North Sydney, NSW (1989b) Wypych, P. W., Pneumatic Conveying of Bulk Solids, PhD Thesis, Department of Mechanical Engineering, Univ. of Wollongong (1989b) Wypych, P. W., “Pressure Drop in Cement Pneumatic Conveying Systems,” Powder & Bulk Solids Conf., Reed Exhibition Companies, pp. 467–471, Rosemont, Illinois, USA (1992) Wypych, P. W., “Optimising & Uprating Pneumatic Transport Systems,” Nat. Conf on Bulk Materials Handling, Yeppoon, Qld., IEAust. Proc., pp. 197–203 (1993) Wypych, P. W., “Latest Developments in the Pneumatic Pipeline Transport of Bulk Solids,” 5th Int. Conf. on Bulk Materials Storage, Handling and Transportation, Newcastle, IEAust, Proc, 1:47–56 (1995a) Wypych, P. W., “Engineering Design in the Pneumatic Pipeline Transport of Bulk Solids,” Mech. Eng. Trans., IEAust, ME20(4):293–298 (1995b) Wypych, P. W., and Arnold, P. C., “The Use of Powder and Pipe Properties in the Prediction of Dense-Phase Pneumatic Transport Behavior,” Pneumatech 2 , Canterbury, England, Organised by the Powder Advisory Centre, London (1984) Wypych, P. W., and Arnold, P. C., “Predicting and Improving Flow Performance in Dense-Phase Pneumatic Transportation,” Int. Symp. on the Reliable Flow of Particulate Solids, Organised by the Chr Michelsen Institute, Bergen, Norway (1985a) Wypych, P. W., and Arnold, P. C., “A Standardised-Test Procedure for Pneumatic Conveying Design,” Bulk Solids Handling, 5(4):755–763 (1986b)

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772 Fluidization, Solids Handling, and Processing Wypych, P. W., and Arnold, P. C., “Pneumatic Transportation and Fluidization Performance of Power Station Fly Ash,” Bulk Solids Handling, 6(1):93–97 (1986a) Wypych, P. W., and Arnold, P. C., “Feasibility and Efficiency of Dense-Phase Pneumatic Transportation,” Mech. Eng. Trans., IEAust, ME11(1):1–5 (1986b) Wypych, P. W., and Arnold, P. C., “Minimising Wear & Particle Damage in Pneumatic Conveying Systems,” Powder Handling & Processing, 5(2):129– 134 (1993) Wypych, P. W., and Hauser, G., “Design Considerations for Low-Velocity Conveying Systems & Pipelines,” Pneumatech 4, Int. Conf. on Pneumatic Conveying Technology, pp. 241–260, Glasgow, Scotland, Powder Advisory Centre, UK (1990) Wypych, P. W., Kennedy, O. C., and Arnold, P. C., “The Future Potential of Pneumatically Conveying Coal through Pipelines,” Bulk Solids Handling, 10(4):421–427 (1990) Wypych, P. W., and Pan, R., “Determination of Air-Only Pressure Drop in Pneumatic Conveying Systems,” Powder Handling & Processing, 3(4):303-309 (1991) Wypych, P. W., and Pan, R., “Pressure Drop due to Solids-Air Flow in Straight Pipes and Bends,” Freight Pipelines, (G. F. Round, ed.), pp. 49-67, Elsevier Science Publishers BV, Amsterdam, Netherlands (1993) Wypych, P. W., and Reed, A. R., “The Advantages of Stepping Pipelines for the Pneumatic Transport of Bulk Solids,” Powder Handling and Processing, Vol 2(3):217–221 (1990) Yang, W. -C., “A Criterion for Fast Fluidization,” Pneumotransport 3, Paper E5, Univ. of Bath, England (1976) Zenz, F. A., “Conveyability of Materials of Mixed Particle Size,” Ind. Eng. Chem. Fund, 3(1):65–75 (1964) Zenz, F. A., “Pneumatic Conveying from Grains to Powders,” Pneumatech 2 , Univ. of Kent, Canterbury, England, Organised by the Powder Advisory Centre, London (1984)

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12 Cyclone Design Frederick A. Zenz

1.0

INTRODUCTION

A cyclone represents the simplest and most easily fabricated piece of processing equipment. In principle it operates as a fluid driven centrifuge in which dispersed matter is condensed into its bulk state against the internal walls of a cylindrical vessel by an accelerated settling rate induced under an enormously higher gravitational field. The phenomenon is identical (Zenz, 1989, Ch. 1) to the condensation of molecular matter from a dispersed vapor state. Because of its simple mechanical structure and, in the past, its lack of criticality to acceptable plant operation, its design approaches have suffered. The literature varies from highly theoretical analyses of the forces governing aerodynamic particle trajectories to the observed performances of diametrically opposed empirical innovations for collecting, for example, shredded paper. The early design approach (Rosin et al., 1932) remains to date as the only complete and logical basis for a dimensioning procedure yielding consistently verifiable results, but only when concomitant with strict attention paid to inlet and outlet peripherals. Industrial consortia (AIMS, 1989 to date) are engaged in continual experimental investigations of such

773

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774 Fluidization, Solids Handling, and Processing peripherals and their effects on destroying or enhancing the levels of performance predicted for the basic designs. Since several readily available references (Zenz, 1975; Zenz, 1989, Ch. 7), replete with worked examples, detail the basic dimensioning procedure, this chapter will be devoted primarily to emphasizing the precautions, essentials and peripherals within the design steps necessary to achieve a desired performance. Though discussed here in terms of a separation of solid particles from a gas, the relationships apply equally well to the separation of particles from a liquid, liquid droplets from a gas, and gas bubbles from a liquid.

2.0

REQUIRED DESIGN DATA

It should be incumbent on any designer, regardless of the procedure adopted, to document all of the following input data in order to substantiate the derivation of any anticipated or verifiable subsequent performance: (i) The viscosity, density and flow rate of the carrier gas entering the cyclone (ii) The particle size distribution of every species of feed particles in terms of the diameter of their aerodynamically equivalent spheres (iii) The apparent densities of the feed particles (iv) Any non-uniformity in the particle concentration gradient in the inlet stream (v) The rate of feed of the particles, as for example, lbs/unit volume of gas or as lbs/ unit time (vi) The required, or desired, particle collection efficiency (vii) The maximum available or permissible pressure drop (viii) Any known limits on particle impact velocity, surface characteristics, or bulk flow ability which would lead to any intolerable particle attrition, or equipment erosion (ix) Any space limitations on overall equipment height or width and the means of removal of the collected solids as for example through a dipleg pipe to a fluidized bed, through a rotary valve or ejector to a bin or conveying line, or intermittently to a sealed container

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Cyclone Design

775

(x) Any pressure or piping limitations dictating the shape of the cyclone inlet or of the transition ducting between cyclones in series.

3.0

CORRELATING FRACTIONAL COLLECTION EFFICIENCY

The traditional (Rosin, et al., 1932) mechanistic approach equates the time necessary for a particle to settle at a Stokes’ law velocity across the width of a cyclone’s inlet duct, to the available residence time of the carrier gas stream in its number of spiral traverses within the barrel. With reference to Fig. 1, this permits solving for the smallest particle size able to cross the entire width and reach the wall in the available time.

; D

Figure 1. Derivation of Dth .

On the inherent assumptions that: (a) all entering particles are uniformly dispersed over the area of the inlet duct, (b) the barrel is a perfect cylinder with a straight inlet duct entering tangentially, (c) the trajectory of the incoming particles will not result in any impingement on the gas outlet duct, (d) that the velocity of the downwardly spiraling gas

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776 Fluidization, Solids Handling, and Processing remains constant, (e) the cyclone is long enough to exceed the depth to which the downward spiraling gas aspires before reversing its net direction through an inner upward spiral toward the gas outlet, (f) the number of spiral traverses is a function of either the gas inlet velocity or outlet velocity whichever is the greater and (g) the inlet dust concentration is no greater than 1 grain/cu.ft. of gas in order that particle travel to the barrel wall rests solely on the aerodynamics of individual particles, it is possible by trial and error to derive generalized correlations of cyclone performance from experimental data, but only when obtained with equipment and conditions all in conformance with the traditional mechanistic approach and with full recognition of the “required design data.” Unfortunately the literature is rampant with data obtained from cyclones of so wide a variety of deviating dimensional ratios and ill defined particle properties, sizes and loadings, that full agreement between designers, theorists and practitioners has not yet been achieved. However, in practice, the enormous number of cyclones installed in identical industrial process applications in the petroleum and petrochemical industries (Zenz, A.P.I., 1975) throughout the world and spurred by a common interest in economics and environmental regulations, has led to the correlations shown as the solid curves in Figs. 2 and 3. t;- "o?n ~ .1' :p t h

Figure 2. Fractional efficiency curves.

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Cyclone Design

777

Ns

Figure 3. Correlation of barrel traverses.

The broad spectrum of derived fractional efficiency curves in Fig. 2 simply reflects the effects of cyclone designs deviating from consistence with the basic premises and from ill defined particle characteristics. The most common source of apparent discrepancies between actual cyclone performance and the predictions based on Fig. 2 lies in the definition of feed particle size and density. Coulter Counter, laser, screen, etc., analyses of size simply reflect a dimension which may not cavalierly be taken as the diameter of aerodynamically equivalent spheres even if the apparent particle density is absolutely known and even if microscopic examination were to confirm perfect sphericity, unless the particle surface were also perfectly smooth. “Shape factors,” intended to correct such measured dimensions to those of aerodynamically equivalent spheres, incorporate surface texture as well as overall irregularities in shape. The apparent aerodynamically intended Bahco or Sedigraph procedures inherently incorporate an assumed apparent particle density which must be independently determined in order to define the diameters of the aerodynamically equivalent spheres. Sorting out the results of measurements of fixed bed pressure drop, incipient fluidization velocity, bulk density, blow out velocities, settling rates, microscopic examination, free fall velocities, etc., (Zenz, 1996) through established correlations can result in consistent design data well worth the effort not only in increasing the credibility of the design procedure, but in avoiding subsequent litigations.

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Cyclone Design

779

distribution of the feed gas inflow along the length L is considered proportionately constant in all cases, then it is found (Zenz, 1989, Ch. 7) that the average inflow velocity, V, over the surface area of the exiting vortex falls in a narrow range between 4 and 5 ft/sec. Based on the average peel off velocity of 4–5 ft/sec, reducing the gas outlet tube diameter lengthens the natural vortex length, L. If L is extended to the point where it would touch the cyclone cone, its centrifugal force will, in effect, “grind” the downwardly spiralling stream of collected solids more forcefully against the cone wall causing erosion in the area of contact. In low loaded cyclones, such contact will also lead to some degree of reentrainment and hence to reduction in collection efficiency. If the gas outlet tube diameter is reduced in order to increase outlet gas velocity to a value greater than the inlet velocity, and thereby decrease Dth to increase collection efficiency at lower overall pressure drop, then the cyclone must also be lengthened to “contain” L. A conservative design clearance between the periphery of the natural vortex length and the cone wall is suggested as 2" as shown in Fig. 5.

L

-

Q

Figure 5. Natural vortex length.

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CONES, DUST HOPPERS AND EROSION

Cones are attached to cyclone barrels simply as a transition to the diameter of the so-called dipleg pipe necessary to allow exit of the stream of collected solids (Zenz, 1975). The cone is irrelevant to the cyclone’s performance; many installations in specific applications, as well as smaller units operating in parallel in multi-clones, are seldom provided with cones. Ideally, a conventional cone should be tapered to the angle of repose of the bulk collected solids under the centrifugal acceleration field within the cyclone. Such sophistication is generally disregarded and the cones sloped at 60 to 80 degrees from horizontal. The steeper this slope, the less the possibility of erosion within the upper end of the dipleg at the point where the stream of collected solids cascades from the edge of the cone’s apex and impacts the dipleg wall. Three patents (MacLean, et al., 1982) assigned to Texaco filed in the latter half of 1980 resulted from an in-house review of all observed instances of erosion of cyclones in Texaco’s FCC units. The patents specify that to minimize erosion and maximize collection efficiency, the length of the cyclone relative to its barrel diameter is to conform to the relationship: H/D = 4.49 - 1.09 (Ao/Ai) However, the eroded regions which led to this relationship were all located in the so-called dust hoppers and not in the cyclone cones. Texaco’s personnel failed to realize that the erosion in the dust hoppers is due to impingement by the stream leaving the edge of the cone’s apex. The simple solution to such erosion is to eliminate the dust hoppers. There are numerous cyclones without dust hoppers operating in the process industries, including FCC installations, which are free of erosion in their cones and dipleg pipes provided that they “contain” the natural vortex length free of touching the cone wall. It is, however, imperative that in fabrication a cyclone cone always be rolled as opposed to being formed on a brake. As illustrated in plan and isometric views in Fig. 6, the condensed stream of collected solids spiralling down on the surface of the cone impinges the metal past any bend (resulting from fabrication on a brake) at an angle of attack steeper than 0 degrees. This has been evidenced in numerous instances as a spiral pattern of eroded holes and parallels the

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Cyclone Design

781

mechanism which leads to dust hopper erosion patterns. Poor quality interior finish in the form of seam weld burrs or lining irregularities induce similar points of severe erosion.

~~

Figure 6. Cones, dust hoppers and erosion.

7.0

CYCLONE INLET AND OUTLET CONFIGURATIONS

Cyclone inlets should always be rectangular in shape rather than round. A round pipe entering a cylindrical barrel affords only one point of tangency, whereas a rectangular duct is tangent along an entire side providing proximity to the barrel wall to a greater number of entering particles. When a high velocity gas stream bearing particulates or liquid droplets enters tangentially into a vertical pipe, the particles or droplets will collect into two streams, one spiralling downwardly and the other

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782 Fluidization, Solids Handling, and Processing spiralling upwardly; the latter eventually reaches a limiting height whereupon gravity causes it to head back down as depicted in Fig. 7. Any domed head (designed to withstand high internal pressure) on a cyclone, therefore, permits such an upward spiral to follow an inward path directly to the top of the gas discharge tube, as in Fig. 8, on which it may then spiral down and be sucked into the exiting gas vortex when it reaches the bottom edge of the gas discharge tube. Inserting a false roof into such an existing design will result in substantial improvement in performance. This situation is equally acute in instances where under high temperature operations a protruding internal gas outlet tube would fail structurally. In such instance, external cooling by forced or natural convection can be provided via a folded annulus as illustrated in Fig. 9. Ideally, the gas outlet tube should protrude below the cyclone roof a distance equal to the height of the inlet duct to minimize bypassing of a portion of the inlet flow directly into the gas outlet. The effect on collection efficiency of shortening the gas discharge tube has been treated in the literature (TerLinden, 1949). The argument that lengthening the tube beyond the inlet duct height will afford a greater gas residence time (greater number of spiral traverses) has never been demonstrated. Since the natural vortex length is measured from the bottom edge of the gas outlet tube, its lengthening must be accompanied with a commensurate increase in overall cyclone length.

Figure 7. Natural split in particle paths.

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Cyclone Design

783

Figure 8. Cyclone roof design.

Figure 9. Dimpled outlet tube.

All cyclone design procedures bear the inherent condition that the solids enter uniformly dispersed over the cyclone’s inlet area. This is generally the case where a cyclone is hung inside a vessel above a

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784 Fluidization, Solids Handling, and Processing contained fluidized bed or where the cyclone is external but close-coupled to the wall of a large vessel as in the case of fluidized bed combustors. If the entering stream is not uniformly dispersed, as induced in Fig. 10 by a preceding elbow, collection efficiency will be poorest in arrangement Fig. 10(a) and exhibit a decrease in efficiency with increase in either loading or inlet velocity. This is attributable to the more densely concentrated stream of solids entering adjacent to the gas discharge tube. The effect is more pronounced with a preceding downflow elbow, as in Fig. 10(a), than with an upflow elbow as in Fig. 10(b), presumably because in Fig. 10(b) there is an outlet tube length over which some particles may have a greater chance to reach the barrel wall rather than all hug the gas discharge tube and escape under its bottom edge drawn in with the exiting gas spiral.

Figure 10. Effect of preceding elbows.

Three generic forms of cyclone can be characterized by their inlet configurations as illustrated in Fig. 11. At equal gas inlet or gas outlet velocities, the tangential generally yields the highest collection efficiency. At very high inlet particle loadings, differences between the three forms

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Cyclone Design

785

may be minimal and the volute the better choice. For example, in a 3-stage series arrangement one could expect to find either of the three as the first stage, possibly a volute for the second stage, but certainly never any other than a tangential for the third stage handling the lowest inlet loading and smallest particle sizes. In general, the greater the extent to which the volute or scroll encircles the barrel, as in Fig. 11(b), the greater the loss in performance. This has been attributed to a disturbance of the spiralling flow paths due to the ledge of the volute inside the barrel (particularly in instances where the ledge has been seen to accumulate a layer of particles impinged upon by succeeding inflowing particles). The generally poorer efficiency of the axial or radial vane inlet of Fig. 11(c) has been correlated as simply reflected in a halving of the effective number of traverses of the barrel in the calculation of its characteristic centrifuged diameter Dth.

/ ,/

Figure 11. Generic cyclone forms.

8.0

THE COUPLING EFFECT

In a single stage installation, overall collection efficiency is not influenced by the external terminus of the gas discharge tube, whether a round elbow, Fig. 12(a), a mitered elbow, Fig. 12(b), or a volute housing,

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786 Fluidization, Solids Handling, and Processing Fig. 12(c). However, in a multistage installation where the gas and uncollected solids from a prior stage immediately enter a close-coupled subsequent stage, the overall 2-stage efficiency can exceed that based on the combined efficiencies calculated for each individual cyclone, if the terminus of the gas discharge tube of the preceding stage is designed to maintain the centrifugal discharge flow. Because of the spin in the exiting gas’s natural vortex, which spin continues in its passage up through the gas discharge tube, there can be an appreciable concentration of particles near (if not directly contacting) the inner wall of the gas discharge tube (Silverman, 1950). If the terminus of the gas discharge tube preserves this centrifugal spiralling flow, via a tangential coupling to the subsequent stage, as illustrated in Fig. 12(d), then in effect the gas discharge tube acts as an axial flow pre-separator or low efficiency, single traverse, intermediate cyclone. The particles entering the subsequent stage are then not uniformly dispersed in the entering gas, as is assumed in applying conventional cyclone design procedures and, therefore, these procedures predict a lower efficiency than actually obtainable in practice. The added efficiency, induced by the ducting arrangement in Fig. 12(d), is referred to as giving rise to the coupling effect.

Figyre

12.

Figure 12. Gas outlet tube design.

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Cyclone Design 9.0

787

PRESSURE DROP

The addition of small amounts of solids (Yuu, et al., 1978; Kane, et al., 1973; Kang, et al., 1989; Knowlton, et al., 1977; Baskakov, et al., 1980) in the gas entering a cyclone at first reduces the pressure drop until, at higher loadings, it increases the pressure drop. The former is attributed to the reduction of viscous drag at the walls (or the “lubrication” by the “rolling” of the particles) while the latter is attributed to increasing pressure drop required to accelerate the greater masses. The equations given in various references (Zenz, A.P.I., 1975; Zenz, 1989, Ch. 7) apply to conditions where the solids loadings exceed that at which any reduction in loading would cause pressure drop to be greater than that given by those equations. The oft-quoted (Stairmand, 1951) equation for cyclone pressure drop applies solely to gas alone bearing no solids. The data shown in Fig. 13 taken from a recent paper (Dry, et al., 1993) are typical of the effect of low loading where the pressure drop at zero solids loading decreases gradually to about 0.4 times this pressure drop, before beginning to increase again by virtue of further increase in loading.

Figure 13. Pressure drop and loading.

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788 Fluidization, Solids Handling, and Processing 10.0 SPECIAL CASES There are instances in specific industries, or in unusual applications, which would appear to exhibit major deviations from all the foregoing “best design practice.” This is most frequently associated with unconventional characteristics of the solids being handled. Figure 14 illustrates a particular case in point. This unit has: (a) essentially no barrel beyond the inlet height, (b) a gas discharge tube with half of its perimeter cut away over about half its length (this portion covered with a coarse screen), (c) a relatively high ratio of inlet height to inlet width (80"/18" = 4.44 as compared to the conventional ~2.2–2.7), and (d) a relatively high outlet to inlet area ratio of 3.14 (compared to 0.4 to 1.5). This unit operates with air under a positive inlet pressure and discharges to ambient. It is fed shredded paper ranging in size from 1/2" × 1/2" to 1/2" × 6". It constitutes essentially a single traverse cyclone, which is sufficient for such large “particles.” The screen on the gas discharge tube allows escape of the air on the inside radius of the first turn and thus “shrinks” the usual natural vortex length to zero. The 24" I.D. cone apex is open to an unsealed collection or baling bin and hence some amount of gas underflow augments discharge, rather than reentrainment or hang-up. The so-called “tangentials” offered for example by American Air Filter Corp. illustrated in Fig. 15 constitutes, in effect, a horizontally oriented version of Fig. 14. The European patent office recently granted two applications to the A. Ahlstrom Corp. (Hyppanen, et al., 1991) relating to what is described in the text, and in the apparatus and method claims, as a centrifugal separator of “square cross-section.” The patent applications make no mention, nor give any examples, of what separation efficiencies can be anticipated. The suggested advantages lie in the fabrication convenience and economy afforded by the flat, as opposed to conventionally curved, water walls. The efficacy of Ahlstrom’s separator lies in the solids, at high loadings, piling up in the corners so that the effective inside diameter approaches the contours of a cylinder. This follows the principle of using TEEs or mitered elbows in pneumatic conveying lines as introduced over 50 years ago and as employed in the separator developed for their TRC process jointly by Stone & Webster Engineering Corp. (Gartside, 1985) and Gulf Oil (now Chevron). That separator, illustrated in Fig. 16, was essentially a square section box, all dimensional design ratios of which were optimized in a modeling program at PEMM-Corp studying 32

12-Oct-2001 JMR

Cyclone Design

789

variations. The optimum configuration yielded 95% collection efficiency of a high fines content FCC catalyst. The arrangement in Fig. 16 represents a half-turn cyclone with restricted drawoff maintaining thereby a dense phase layer of collected solids as a contoured erosion shield. Figure 17 illustrates another form of a “less than one turn” cyclone in which the rectangular box of Fig. 16 is now an open channel allowing even more rapid separation of gas and solids. This arrangement has achieved a factor of three reduction in catalyst losses in several operating FCC regenerators. FCC catalyst separation in the developer’s (Zenz, 1988) laboratory tests yielded collection efficiencies greater than 99%. The important aspect of the curving channels in Fig. 17 is that they maintain the collected solids as a dense phase stream at all times as they terminate tangential with a collection vessel’s wall. I oE

I. -~I

I,~

..

I

I

"" I

\

I '1 /1 ~...

,1.'

/

I

Figure 14. Shredded paper collector.

12-Oct-2001 JMR


Figure 15. Tangential single turn cyclones.

~

Figure 16. Rectangular mitered elbow cyclone.

Cyclones are still simply gravity settling chambers in which a particle’s settling velocity is increased by making the particle heavier (through a centrifugal field greater than 32.2 ft/sec/sec and in which sufficient residence time is provided via an induced number of barrel traverses to achieve settling over a distance Lw with the predictability of performance dependent on attention to all design peripherals.

12-Oct-2001 JMR

Cyclone Design

791

Figure 17. Curved “ARM” collector.

11.0 BED PARTICLE SIZE DISTRIBUTION AND CYCLONE DESIGN In a fluidized bed reactor, entrained particles leaving in a dilute phase stream are conventionally and desirably either partially or wholly condensed into a bulk stream and returned to the bed via a centrifugally driven cyclone system. At equilibrium, or when steady state operation is attained, any particle loss rate from the cyclones, as well as the remaining bed particle size distribution, are functions of (a) the rate of any particle attrition within the system and (b) the smallest particle size that the cyclone system was designed to completely collect (i.e., with 100% efficiency), or conversely the largest size which the system cannot recover. These two functions result in an interdependency between loss rate and bed particle size distribution, eventually leading to an equilibrium state (Zenz & Smith, 1972; Zenz, 1981; Zenz & Kelleher, 1980). Case A. To illustrate, consider a situation in which a reactor is charged with a catalyst having a size distribution ranging from 1 to 150

12-Oct-2001 JMR

792 Fluidization, Solids Handling, and Processing microns. Suppose that the cyclone system was designed to retain all particles 10 microns or larger in size and that the particles are so “strong” that there will be no attrition. As time progresses after start-up, the bed will approach a size distribution from 10 to 150 microns. All particles smaller than 10 microns will have been lost, only the fluidizing off gases will leave through the cyclone system. No additional catalyst will need to be charged to make up for the loss from the cyclone system because no further particle loss occurs once the 0 to 10 micron fraction of the charge has left the system. The mean particle size of the equilibrium bed will be larger than that of the original catalyst charge, since the weight percent of all sizes smaller than the largest will have decreased in proportion to the lost weight fraction of the 0 to 10 micron particles in the original charge. Case B. Suppose, more realistically, that the catalyst undergoes a known, experimentally determined, rate of attrition as a function of particle size (Zenz, 1971; Zenz & Kelleher, 1980). The particle loss rate from the cyclone system will now approach and finally equal the rate of production of 0 to 10 micron particles by attrition from all the larger sizes. To maintain reactor inventory, this loss rate will be replaced, at an equal rate, with fresh catalyst. Since the rate of attrition of any size particle depends on its concentration in the stream subjected to the attrition (as finer particles effectively “cushion” the coarser), and since the loss is replaced with fresh catalyst (containing the coarsest), the bed size distribution will reach a steady state between 10 and 150 microns in which the mean size, as well as all sizes smaller than the largest, will now be decreased from what would have prevailed under conditions of zero attrition. Case C. Now suppose that in order to maintain conversion in the reactor the equilibrium bed must be further increased in catalytic surface area, or in other words must maintain a greater “fines” content under steady state operation. This would require the cyclone system to now retain all particles smaller than say 5 microns in diameter. Since the existing system design could not retain particles smaller than 10 microns, the cyclone system must be modified to meet the new and more stringent specification. This can be accomplished either by replacing the final collection stage with a larger number of smaller cyclones in parallel, still operating at the same inlet and outlet velocities, or by modifying the existing cyclone system’s final stage to operate at higher inlet and outlet

12-Oct-2001 JMR

Cyclone Design

793

velocities. Both changes are designed in principle to reduce the time required for particles to reach and “condense” on the barrel walls. Replacing the final stage with a multiplicity of smaller units is likely the more costly modification, but would advantageously not increase the catalyst attrition rate. The less costly expedient of simply modifying the internal dimensions of the inlets and outlets of the existing cyclones would simultaneously increase the catalyst attrition rate, so that when the new steady state is reached, the bed would have a size distribution from 5 to 150 microns, a mean particle size smaller than in all previous cases, a greater weight fraction of all sizes smaller than the largest, and a loss rate of 0 to 5 micron particles likely greater than the weight rate of loss of 0 to 10 micron particles in case B. This will require more fresh catalyst makeup to maintain reactor inventory. The time required to reach the new steady state equilibrium will be principally a function of the entrainment rate and the reactor catalyst inventory. There are cases illustrated in the literature in which this could take as long as a year (Zenz & Smith, 1972). The choice of cyclone modification, from an operating point of view, becomes a balance of incremental profit from increased conversion, versus catalyst makeup charges, and from a capital cost point of view, the price of either of the cyclone modifications, which must be depreciated. In many instances, there is an additional background time element, involving ongoing development of more attrition resistant and/or active catalyst.

12.0 CENTRIFUGAL VERSUS CENTRIPETAL CUT POINT PARTICLE SIZE Though the efficiency of a cyclone is relatable to the smallest particle size it will collect with 100% efficiency (or conversely the largest particle that appears in its loss), it must be borne in mind that any such evaluation, or comparison among designs, must be conducted under identical conditions of loading and feed particle size distribution as well as gas density, viscosity and inlet velocity. In operation a cyclone, losing a substantial amount of fine particles when fed a high concentration of fines is not necessarily to be regarded as an inherently poor design. It should principally be evaluated in terms of the largest particle that it loses (i.e., is unable to collect). In a similar situation in which such a cyclone acted as a second stage where its preceding unit recovered a reasonable amount of

12-Oct-2001 JMR

794 Fluidization, Solids Handling, and Processing small particles or where it was preceded with less attrition, the loading to this now second stage would have a lower percentage content of small particles and hence would show a higher total efficiency. This higher efficiency simply reflects a feed containing less losable material; the maximum size lost would still be the same as when fed material with a greater concentration of the unrecoverable particle sizes, where it had then been considered a poorer efficiency cyclone. The differentiation between collection efficiency and maximum unrecoverable or lost size has never been clearly established. Maximum unrecoverable particle size is overshadowed at high loadings where bypassing to the outlet tube can cause larger particles to be lost though the overall efficiency may be high. Fundamentally the cut point diameter or maximum size lost is observable only at low loadings less than about 50 grains/cu.ft. and reasonably predictable by equating the centrifugal force of the exiting vortex (impeding particle loss) to the net or average radial inflow (centripetal force) to the exiting vortex. The low loading (exiting) cut point diameter differs in its derivation and hence its significance from the higher loading (“inlet catch”) Dth of Figs. 1 and 2. A lack of appreciation between these two distinct, albeit seemingly interrelated, performance criteria is likely one of the causes of the 5-fold variance in the intercepts of experimental fractional efficiency curves, at 100% collection, illustrated in Fig. 2.

13.0 CYCLONE DESIGN EXAMPLES With consideration to the domestic industrial practitioner, the following illustrative examples utilize conventionally recognizable units as viscosity in centipoise, density in lbs/ft3, velocity in ft/sec, vessel dimensions in inches, and particle diameter in microns. Example A. Suppose it was desired to design a cyclone to collect all particles larger than 97 microns and release no more than 60 lbs/hr from a 6328 ACFM gas stream bearing 80 grains of solids per cubic foot of gas. The particle size distribution and density of the solids is given in Fig. 18; the gas has a density of 0.1 lbs/ft3 and a viscosity of 0.02 centipoise. The pressure drop across the cyclone is not to exceed 10" H 2O. Note from Fig. 1 of the text that 100% collection efficiency of 97 micron particles would require a Dth of 9.7 microns which would be collected at an average 50%

12-Oct-2001 JMR

Cyclone Design

795

efficiency. From considerations of cost, capacity, size and inlet pressure drop, the optimum ratio of height-to-width of a rectangular inlet lies between 2 and 3 (avg. = 2.5) so that cyclone capacity can be expressed as:

Eq. (1)

ACFM = 2.5L2wV1n 60

Solving for Lw and substituting in the definition of Dth given in Fig. 1 of the text: 9.7 9 × 0. 02 × 0. 000672 6,328 = 12 × 25,400 3. 14N sV1 (85 − 0. 1) n 150V1

which for the case of a single unit (i.e., n = 1) reduces to Eq. (2)

Ns V11 .5 = 2908

Figure 18. Particle size analysis of cyclone feed.

12-Oct-2001 JMR

796 Fluidization, Solids Handling, and Processing From Fig. 3 of the text, a velocity of 75 ft/sec and its corresponding NS of 4.5 would satisfy (2), in which case, from Eq. (1): LW = 12

6328 = 9 inches 2. 5 × 75x60

therefore, inlet height = 22.5", and from Fig. 19, gas outlet tube I.D. = 9 × 2 = 18"; assuming 1/4" wall thickness, gas outlet tube O.D. = 18.5". Therefore, barrel I.D. = 2(9) + 18.5 = 36.5" and from Fig. 19, Lmin = 54". Minimum dipleg I.D., d, conservatively assuming 100% collection efficiency and 20% safety factor is calculable from the dimensionless gravity flow equation:

6328 × 80 × 4 × 144 = 1.2 × 43 d ½ 60 × 7000 d 2 3 .14 from which d = 1.8" or a design minimum of 1.2 × 1.8 = 2.2" which from Fig. 19 calls for a 4" industrial minimum, and thus in summary, a cyclone having the dimensions shown in Fig. 20. The overall collection efficiency shown in Table 1 is determined from Fig. 2 of the text for the feed distribution given in Fig. 18. From Fig. 4 of the text at 80 grains/cu.ft., EL = 98.7%, therefore: Loss Rate =

6328 × 60 × 80 (1 − 0 .987 ) = 56.4 lbs/hr 7000

which meets specifications on performance (i.e., less than 60 lbs/hr loss rate). The anticipated pressure drop is calculable from the relationships in Fig. 21. Assuming the cyclone is located within, or attached externally to, the shell of an 8' diameter fluidized bed reactor operating at a superficial gas velocity of 2 ft/sec, then:

(22.5 × 9 )4

Area ratio = 144(8)2 3.14 = 0.028; hence K = 0.5

12-Oct-2001 JMR

798 Fluidization, Solids Handling, and Processing Table 1. Prediction of Cyclone Collection Efficiency

Inlet contraction loss: ∆P(s-i)G = 0.00298(0.1)[1.5(75) 2-(2) 2] = 2.513" H2O Solids acceleration loss: ∆P(s-i)G = 80(75) 2/7000(167) = 0.385" H2O Barrel loss: d hi = 4(22.5 × 9)/2(22.5 + 9) = 12.857"

12-Oct-2001 JMR

Cyclone Design

--'r-

Homen!:l.ture .P

VG.

.Pressure Drop, In. H.O ..Load1n9 Lb So11ds/Ft. ot aas Veloc1ty, Ft/Sec. .Gas Dens1ty, Lb/Ft. .Gas V1scos1ty. Lb/r~ Sec. .D1&meter

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Figure 21. Pressure drop relationships.

12-Oct-2001 JMR


Re =

12.857 × 75 × 0.1 = 597,888 12 × 0. 02 × 0.000672

from any Fanning friction factor chart for smooth drawn tubing, f ~ 0.0033 ∆PB = 0.0033(36.5)(0.1)(4.5)(75) 2/26.2(12.857) = 0.905" H2O Reversal loss: ∆Pr = 0.1 (75) 2/335 = 1.679 “ H2O Exit contraction loss: Ve = 6328(4)(144)/3.14(60)(18) 2 = 59.68 ft/sec VB = 6328(4)(144)/3.14(60)36.52 = 14.5 ft/sec Area ratio = (18/36.5) 2 = 0.243; hence K = 0.415 ∆P e = 0.00298(0.1)[1.415(59.68) 2-(14.5) 2 ] = 1.439" H2O Total calculated ∆P across cyclone = 2.513 + 0.385 + 0.905 + 1.679 + 1.439 = 6.921" H2O if specific solids inlet loading were >10 lbs/sec × ft2.

Specific loading =

6328 × 80 × 144 = 0.857 lbs/sec x ft2 60 × 7000 × 22.5 × 9

From Fig. 21 at an abscissa of 0.01 + 0.857 = 0.867, the ordinate equals 1.32; therefore:

12-Oct-2001 JMR

Cyclone Design

801

True cyclone ∆P = 1.32 × 6.921 = 9.136" H2O, meeting specifications Example B. Suppose in the previous example that the loss rate had been specified as not to exceed 40 lbs/hr. To minimize the increase in pressure drop accompanying any increase in inlet velocity necessary to reduce Dth to a value which would bring performance up to the desired level, it might be more expedient to instead increase the gas discharge velocity. To find EL required to meet 40 lbs/hr loss rate set: 40 = (1 - EL)6328(60)(80)/7000; from which EL = 99%. From text Fig. 4 at 80 grains/ft3, required EO = 94%. The correspondingly required Dth is estimated by ratio from Example A: Dth 1 − 0.94 ~ 9.7 1 − 0.9086

from which Dth = 6.3 microns (estimated) and now checked against the correlation in Fig. 2 of the text for the feed distribution in Fig. 18 (see Table 2) The 95.41% is slightly better than the required 94% so that the foregoing could be repeated with a new estimate of Dth slightly greater than 6.3 microns. However, since the difference between 94% and 95.41% is probably within the accuracy of any practical means of measurement, the value of 6.3 microns will be accepted as the design criterion, in which case, from Fig. 4, EL ~99.14% leading to a predicted loss rate of: (1 - 0.9914)6328(60)(80)/7000 = 37.3 lbs/hr which meets the specified 40 lbs/hr limit. From the definition of Dth in Fig. 1 of the text: 6. 3 9( 0.02)0.000672(9) = 12(25, 400) 3. 14Ns V (85 − 0. 1)12

12-Oct-2001 JMR

802 Fluidization, Solids Handling, and Processing or NS Ve = 796.15 which from Fig. 3 suggests Ve = 140 ft/sec and NS = 5.7, so that DO ~11.75" and from Fig. 19, L = 82", which results in a unit dimensioned as shown in Fig. 22.

Table 2. Prediction of Collection Efficiency for Cyclone in Example B

12-Oct-2001 JMR

Cyclone Design

803

Figure 22. Cyclone design dimensions for Example B.

The cyclone in Fig. 22 will have a greater contraction loss than that in Fig. 20: Area ratio = (11.75/36.5) 2 = 0.1036; hence K ~ 0.47 ∆Pe = 0.00298(0.1)[1.47(140) 2 - (14.5) 2] = 8.52" H2O ∆P (calc’d) = 2.513 + 0.385 + 0.905 + 1.679 + 8.52 = 14" H2O

12-Oct-2001 JMR

804 Fluidization, Solids Handling, and Processing From Fig. 21, at a solids inlet rate of 0.857 lbs/sec·ft2, the true ∆P/14 = 1.32, so true ∆P = 18.48" H2O which exceeds the specified maximum of 10" H2O.

14.0 ALTERNATE APPROACH TO SOLVING EXAMPLE B If, as an alternative, the Dth of 6.3 were to be achieved by increasing the inlet velocity, then Eq. (2) of Example A becomes NS (Vi )1.5 = 6895 which from Fig. 3 suggests Vi = 117.7 ft/sec and NS = 5.4,

then:

LW = 12

ACFM = 7 3 /16 " 150 V1

therefore: inlet height = 2.5(7 3 /16 ) = 18"; DO = 14 3 /8 "; DB = 29.25" and L = 67" which results in a cyclone dimensioned as shown in Fig. 23. Pressure drop through the cyclone in Fig. 23 is again calculated from the relationships in Fig. 21. As in Example A,

Area ratio =

18 × 7. 1895× 4 3.14(8) 144 2

= 0.0179; K = 0.5

Inlet contraction loss: ∆P(s-i)G = 0.00298(0.1)[1.5(117.4) 2 - (2) 2] = 6.160" H2O Solids acceleration loss: ∆P(s-i)P = 80(117.4) 2/7000 × 167 = 0.943" H2O

12-Oct-2001 JMR

Cyclone Design

805

Figure 23. Alternative cyclone design for Example B.

Barrel loss: d hi = 4(18 × 7.1875)/2(18 + 7.1875) = 10.27"

Re =

10.27 × 117.4 × 0.1 = 747,796; f = 0.00314 12 × 0.02 × 0.000672

∆PB =0.00314(29.25)(0.1)(5.4)(117.4) 2/26.2(10.27) = 2.54" H2O

12-Oct-2001 JMR

806 Fluidization, Solids Handling, and Processing Reversal loss: ∆Pr = 0.1(117.4) 2/335 = 4.114“ H2O Exit contraction loss: Ve = 6328(4)144/60(3.14)(14.375) 2 = 93.58 ft/sec VB = 6328(4)144/60(3.14)(29.25) 2 = 22.60 ft/sec Area ratio = (14.375/29.25) 2 = 0.2415; K = 0.415 ∆Pe = 0.00298(0.1)[1.415(93.58) 2-(22.6) 2] = 3.54" H2O ∆P (calc’d) = 6.16 + 0.943 + 2.54 + 4.114 + 3.54 = 17.3" H2O Inlet ldg. = 6328(80)144/60(7000)18(7.1875) =1.342 lbs/sec·ft2 Therefore from Fig. 21 True ∆P = 1.26 × 17.3 = 21.8" H2O This again exceeds the specified 10" H2O maximum to a greater degree than the cyclone in Fig. 22, although it represents also a substantially smaller unit. Since neither design in Figs. 22 or 23 will satisfy the specified pressure drop limitation, the only alternative lies in reducing inlet width without increasing inlet velocity, which requires multiple cyclones operating in parallel as illustrated in the design of Example C. Example C. Suppose, in the previous Example B, performance must be achieved at a pressure loss less than 10" H2O requiring therefore multiple cyclones in parallel.

12-Oct-2001 JMR

Cyclone Design

807

At Dth = 6.3, Eq. (2) of Example A now reduces to NS (Vi)1.5 n 0.5 = 6895 If n = two cyclones in parallel, then NS (Vi)1.5 = 4876 which, from Fig. 3 of the text, would be satisfied with Vi = 97.3 ft/sec and NS = 5.08 so that LW = 12(6328/150(97.3)2) 0.5 = 5-5/8" (to nearest 1/16") Therefore at Vi = 97.3 ft/sec and inlet height = 13-7/8" DO = 11.25"; DB = 23", and L = 43" which results in 2 cyclones each dimensioned as shown in Fig. 24.

Figure 24. Cyclone design for Example C (one of two parallel units).

12-Oct-2001 JMR

808 Fluidization, Solids Handling, and Processing The pressure drop through 2 parallel units of Fig. 24 at Vi = 97.3 ft/ sec, NS = 5.08, Ve = 76.4 ft/sec, VB = 18.28 ft/sec, with inlet area ratio of 0.0108 so that K = 0.5, from Fig. 21: Inlet contraction loss: ∆P(s-i)G = 0.00298(0.1)[1.5(97.3) 2-(2) 2] = 4.23" H2O Solids acceleration loss: ∆P(s-i)P = 80(97.3) 2/7000(167) = 0.648" H2O Barrel loss: d hi = 4(5.625)(13.875)/2(5.625+13.875) = 8" Re = 8(97.3)0.1/12(0.02)0.000672 = 482,639; f = 0.00347 ∆PB = 0.00347(23)0.1(5.08)(97.3) 2/8(26.2) = 1.831" H2O Reversal loss: ∆Pr = 0.1(97.3) 2/335 = 2.826" H2O Exit contraction loss: Area ratio = (11.25/23) 2 = 0.239; K = 0.415 ∆Pe = 0.00298(0.1)[1.415(76.4) 2-(18.28) 2] = 2.362" H2O ∆P(calc’d) + 4.23 + 0.648 + 1.831 + 2.826 + 2.362 = 11.897" H2O Inlet ldg.= 6328(80)144/2(60)7000(5.625)13.875 = 1.11 lbs/sec·ft2

12-Oct-2001 JMR

Cyclone Design

809

From graph in Fig. 21: True ∆P = 1.29 × 11.875 = 15.35" H2O This still exceeds the design specification of 10" H2O and could have been anticipated, since to meet this pressure drop criterion, Vi from Example A must not exceed about 75 ft/sec which would require 6 cyclones in parallel as opposed to the 2 in this Example C.

15.0 ALTERNATE APPROACH TO SOLVING EXAMPLE C Assume the designer does not desire to use 6 cyclones in parallel, but must still meet all the specifications in Example B (i.e., cannot accept 15.35" H2O pressure drop) but is willing to accept 4 cyclones in parallel. Could 4 cyclones suffice (by a reduction in inlet velocity compensated in performance by an increase in exit gas velocity)? Since true ∆P must be ≤ 10" H2O, then ∆P(calc’d) must equal or be < ~10/1.3 = 7.7" H2O. In Example C, ∆P(calc’d) exclusive of exit loss equaled 9.535" H2O so that now the following 3 equations would need to be satisfied with n = 4: Eq. (3)

9.535(Vi /97.3) 2 + (Ve /76.4) 2 = < 7.7

Eq. (4)

LW = 12[6328/60(2.5)Vin]0.5

Eq. (5)

6.3/12(25400) = [9(0.02)0.000672LW /3.14(12)NSVe(85-0.01)12] 0 . 5

Equations (3), (4) and (5) reduce to: Eq. (6)

Vi2 + 0.4018Ve2 = 7645.3

Eq. (7)

LW = 38.97/(Vi)0.5

Eq. (8)

NS Ve (Vi)0.5 = 3447.3

12-Oct-2001 JMR

810 Fluidization, Solids Handling, and Processing Since NS is a function of Ve , solving Eq. (8) for Vi and substituting in Eq. (6) leads to Vi = (3447.3/NSVe)2 which is satisfied with Ve = 87 ft/sec at NS = 4.81 so that Vi = 67.86 ft/sec LW = 4.75" DO = 7.5" DB = 19.5" L = 32" inlet height = 11.75" as summarized in Fig. 25. Check on performance: Dth = 12(25400)[9(0.02)0.000672(4.75)/12(3.14)87(85-0.1)4.81] 0.5 = 6.3 microns (will meet specification) Check on pressure drop: Inlet contraction loss: ∆P(s-i)G = 0.00298(0.1)[1.5(67.86) 2-(2) 2] = 2.057" H2O Solids acceleration loss: ∆P(s-i)P = 80(67.86) 2/7000(167) = 0.315" H2O

12-Oct-2001 JMR

Cyclone Design

811

Barrel loss: d hi = 4(4.75 × 11.75)/2(4.75 + 11.75) = 6.77" Re = 6.77(67.86)0.1/12(0.02)0.000672 = 284,854; f = 0.0039 ∆PB = 0.0039(19.5)0.1(67.86) 24.81/26.2(6.77) = 0.95" H2O Reversal loss: ∆Pr = 0.1(67.86) 2/335 = 1.375" H2O Exit contraction loss: Area ratio = (7.5/19.5) 2 = 0.148; K = 1.45 VB = 6328(4)144/4(3.14)60(19.5) 2 = 12.7 ft/sec Pe = 0.00298(0.1)[1.45(87) 2-(12.7) 2] = 3.223" H2O ∆P(calc’d) = 2.057 + 0.315 + 0.95 + 1.375 + 3.223 = 7.92" H2O Inlet ldg. = 6328(80)144/4(60)7000(4.75)11.75 = 0.778 lbs/sec·ft2 From the graph in Fig. 21: True ∆P = 1.33(7.92) = 10.53" H2O This barely exceeds the specified 10" H2O and can be brought to specification by a moderate increase in inlet height to 12" which reduces Vi to 66.6 ft/sec and results in the overall dimensions shown in Fig. 25. The foregoing examples illustrate the relationships among the variables as they affect performance (collection efficiency) and pressure drop.

12-Oct-2001 JMR


5-3/4

II

Figure 25. Alternative cyclone design for Example C (one of four parallel units).

16.0 DIPLEG SIZING AND CYCLONE PRESSURE BALANCE Dipleg Diameter. In the foregoing examples, the minimum 4" I.D. (see Fig. 19) was sufficiently large to pass the collected solids. The basic relationship for sizing diplegs takes the dimensionless form:

( g d /12) ρ 4W 144 = 3. 14d 2 ( tanα )0.5 B 0.5

12-Oct-2001 JMR

ρ B −ρ G ρB

Cyclone Design where:

813

W = lbs of solids collected per second d = dipleg I.D. inches (minimum) α = collected solids’ angle of internal friction, degrees ρ B = collected solids’ bulk density, lbs/cu.ft. ρ G = density of gas entering cyclone, lbs/cu.ft. g = gravitational field, ft/sec2

If the fluid stream is a gas, the last term in the above equation is essentially unity. Unless the cyclone itself is rotating or, for example, located on another planet, g can be taken as 32.2. If the bulk solids’ angle of internal friction is unknown, then taking an average value of 62 degrees, the equation reduces to: d = (152.8 W/ρ B)0.4 W is generally based conservatively on the assumption of 100% collection efficiency. Pressure Balance. The extent to which the dipleg of a cyclone is filled with exiting solids depends on the pressure balance around the cyclone and its dipleg. Figure 26 shows a cyclone with pressure PB (psi) inside its barrel and with its dipleg immersed Hbed feet into a fluidized bed of density ρ bed (lbs/ft3). A simple pressure balance treating the dipleg as one leg of a manometer gives the feet of solids in the dipleg (HDL ) from: PB + (HDL ρ DL )/144 = PV = ( ρ Bed HBed)/144 where ρ DL is the lbs/ft3 bulk density of solids in the dipleg and PV is the pressure in the vessel in the dilute phase above the bed, psi. (PV -PB) is determined from the cyclone pressure drop correlations in Fig. 21.

12-Oct-2001 JMR


Figure 26. Cyclone pressure balance (no restrictive “valve” on dipleg end).

NOTATIONS

12-Oct-2001 JMR

Ai

Gas inlet duct area

Ao

Gas outlet tube area

D

Barrel I.D.

Do Dp

Gas outlet tube I.D. Particle diameter

Dth

Smallest particle able to cross width Lw

EL

Collection efficiency at actual inlet loading

Eo

Collection efficiency at q/m > 1.0 µC/kg

In general, the data of Table 4 tend to confirm this estimated range; however, the upper limit of 1.0 µC/kg is considerably exceeded in the case of data obtained with individual particles of tribo-active, insulating polymers. Whether or not significant volume separation of positive and negative charges occurs with these polymers is not known. Thus, the validity of the upper limit in Eq. (9) is difficult to judge. For fluidized beds consisting of two or more particulate constituents (for example, distinct size cuts or different materials), charge separation is virtually inevitable. An example of bipolar charging is provided by crushed coal which, in the narrow size range from 63 to 75 µm, exhibits a bipolar charge distribution ranging from about -30 to +30 µC/kg with the average charge near zero (Harris, 1973). These values depend strongly on sulfur content. The likely reason for this strong bipolar charging is the heterogeneous nature of coal; when crushed, individual particles vary considerably in their mineral makeup. It should be pointed out that particle-particle triboelectrification does not create significant net volume charge ρ unless the various particle populations segregate within the bed. As shown in Sec. 2.4, only charge segregation can create strong electrostatic fields and the associated high electrostatic potentials. If strong particle-wall triboelectrification occurs in a bed of homogeneous particles, then segregation is much more likely.

Electrostatics and Dust Explosions 835 Methods of Charge Reduction. There are a number of methods for reducing electrostatic charging in fluidized beds, including humidification of the fluidizing air, conductive particle coating, reduced gas flow, careful selection of vessel materials, the choice of more conductive particulate, and finally of course the addition of antistatic agents. As already mentioned, the application of humidified air to reduce electrification in fluidized beds dates from some of the earliest investigations (Osberg and Charlesworth, 1951). Many investigators have studied its practical usefulness (Katz, 1957; Davies and Robinson, 1960; Tardos and Pfeffer, 1980; Guardiola et al., 1996). In general, the conclusion is that relative humidities in the 60% to 70% range are required to reduce charge accumulation significantly. Even then, at least one report observed that the electrostatic effects do not correlate very well with measured specific charge values (Wolny and Kazmeirczak, 1989). There is also the problem that high humidities can increase the cohesive properties of some particles. Like humidification, conductive coatings on particles can reduce electrostatic charge by accelerating the rate at which the charge leaks away from the bed to ground (Katz, 1957; Boland and Geldart, 1972). Unfortunately, many coatings are not sufficiently robust to last very long in a vigorously churning, bubbling bed, especially at elevated temperature. The same problem exists with most of the modern antistatic agents, of which there are many types (anon., 1993). These materials, which work by making the particle surface attract and hold moisture, are sufficiently effective to improve flow characteristics of powders but the coatings cannot be expected to stay intact for long in a fluidized bed. Another method to reduce the effects of charging is to add a small quantity of fines (conductive or nonconductive). For example, graphite, added to beds of glass beads in amounts of ~0.14% by weight, has been found to reduce electrostatic activity (Bafrnec and Beña, 1972). In similar tests, finely ground coal, titanium dioxide, pigment particles, and aluminum powder were introduced into beds of tribo-active polystyrene beads (Wolny and Opalinski, 1983; Wolny and Kazmeirczak, 1989). The mechanism by which the fines reduce tribocharging is believed to be that the smaller particles coat the larger particles and reduce the number of intimate contacts. It is also true that, if the fines are oppositely charged, then the coating will simply shield the particle charge and reduce the effective specific charge q/m. In one interesting experiment, it was shown that the addition of alumina fines to a particle bed of the same material reduces charging, apparently via the same mechanism of the fines coating the larger

836 Fluidization, Solids Handling, and Processing particles (Singh and Hsieh, 1987). This successful charge reduction method is a good example of size-dependent tribocharging. Binary mixtures of glass ballotini and steel beads of nearly equal size (~500 µm and ~275 µm, respectively) also exhibit reduced electrostatic activity, as measured by suspended electrodes (Guardiola et al., 1992). The highly conducting steel beads were added to the glass particles in amounts up to approximately 10% by volume. The obvious measures of lowering superficial gas velocity in the bed and the selection of wall and/or particulate materials certainly exist. But, in a given process, there is often no practical way to implement such solutions. Fluidization conditions are dictated by the application for which the bed is intended; furthermore, the choice of particulate materials can seldom be made on the basis of electrostatic considerations.

3.3

Beneficial Effects of Electric Charge

Though the focus of this chapter is electrostatic nuisances and hazards associated with fluidization, it must be recognized that in some cases particle charging can be exploited with beneficial effect. Applications that have been proposed and investigated include enhanced collection of gas-entrained particulate pollution such as fly ash in fluidized bed combustion gas cleaners. These electrofluidized beds (or EFB’s) function essentially as electric field-assisted gas scrubbers, efficiently removing charged particulate pollution from a gas stream (Melcher et al., 1979a,b). Electrofluidization has also been exploited in powder coating processes (Lupinski and Gorowitz, 1979). A second beneficial electrostatic effect is the electroclamping phenomenon. Here, an imposed DC electric field induces interparticle forces strong enough to suppress bubbling and to control the recirculation of particles in fluidized or spouted beds (Johnson and Melcher, 1975; Talbert et al., 1984; Law et al., 1996). Note that electroclamping is related to strong induced dipole-dipole interactions between particles and does not depend on tribocharging.

4.0

ESD DUST IGNITION HAZARDS

The next task in considering the electrostatic hazards associated with a fluidized bed is to consider the specific conditions required for the ignition of suspended dust and powders. It is crucial to cover this subject because

Electrostatics and Dust Explosions 837 dust explosions occur only under very stringent conditions imposed upon concentration, oxygen content, and ignition source spatial and temporal characteristics.

4.1

Basics of Suspended Solids Ignition

Preliminaries. The combustion of suspended dusts and powders is quite complex and only imperfectly understood. The complexity stems from both fundamental and practical considerations. On the fundamental side, the ignition of suspensions of finely divided solids is influenced by hard-to-quantify factors such as the time-varying concentration of solids, the chemical activity and morphology of the particulate, and the degree of confinement provided by the vessel. On the practical side, industrial conditions are seldom sufficiently well-controlled or characterized to justify application of existing theoretical models. For all the above reasons, this chapter can provide only a very abbreviated coverage of ignition basics. The reader is referred to other sources for in-depth treatment of dust and powder explosions (Bodurtha, 1980; Bartknecht, 1981; Bartknecht, 1987). Fire Triangle. In the context of industrial fire and explosion safety, the universally recognized means to introduce the necessary conditions for fires and explosions is the fire triangle. The fire triangle, shown in Fig. 4, signifies the three essential requirements for a fire or explosion: (i) fuel, (ii) an oxidizing agent, and (iii) an ignition source. The absence of any one of these elements means that no fire or explosion can occur. In dust explosions, the dust is the fuel while the oxidant is almost always atmospheric oxygen. There are certain powder mixtures that can support potentially hazardous exothermic reactions—the thermite reaction in mixtures of cuprous oxide and aluminum powder is one—but these do not fit the conventional model of a dust deflagration and are not of interest here. In addition to electrostatic discharges, many other ignition sources must be acknowledged, including open flames, overheated bearings, and sparks caused by short circuits in electric power equipment. Of all these, ESD is simultaneously the least likely to cause an ignition and the most difficult to control, an irony due to the seemingly capricious nature of electrostatics. The investigation of an industrial fire or explosion invariably starts with an effort to identify each element in the fire triangle. Quite often, the fuel is known and so either the oxidizer or the ignition source becomes the focus of the investigation. For example, in a blender used to mix a

838 Fluidization, Solids Handling, and Processing flammable and strongly triboelectrically active powder, measures might have been taken to inert the vessel, that is, to reduce the oxygen level below the concentration where a fire is possible. If a fire occurs anyway, then the focus of the investigation becomes why and how the inerting system failed. If, on the other hand, measures had been taken to dissipate triboelectric charge, for example, by the addition of an antistatic agent, then suspicion might be cast upon the efficacy of the agent or on the possibility that some other ignition mechanism, such as an overheated bearing, might be responsible.

FUEL

OXIDIZER

IGNITION SOURCE

Figure 4. The well-known fire triangle, which depicts the three essential requirements for fires and explosions.

Concentration Conditions for Ignition. The gaseous O2 and particulate solids concentration conditions required for ignition of dusts and powders suspended in air are far more stringent than those for flammable gases and vapors. It is for just this reason that dust ignition events are in fact relatively rare compared to vapor and gas ignitions. It seems that the probability for simultaneous occurrence of the right concentration values for oxygen and dust, along with a sufficiently energetic ignition source, is low. For a dust ignition to occur, the suspended solids concentration must lie between lower and upper limits which vary from material to material and are influenced by subtle factors such as particle shape and size distribution. From the standpoint of assessing safety hazards in commercial and industrial operations, the lower explosive limit is the more important one. The rationale of this statement is that, if the possibility of exceeding the lower limit in a powder-handling operation can not be completely ruled out, then a hazard must be recognized and appropriate measures taken. Over the

Electrostatics and Dust Explosions 839 years, lower explosive limit data have been obtained for the dusts of many important powders and granular solids (National Fire Protection Association, 1986). Table 5, though only representative, shows that most dusts have a minimum explosion concentration in the range from ~30 to ~100 grams per cubic meter. For a cloud of 10 µm diameter polymer particles, this figure translates to ~10 particles per cubic millimeter. A useful rule of thumb helpful in interpreting this result is that a dust cloud in the explosive range will appear optically opaque. If visibility is heavily obscured by a cloud of suspended dust, then it is best to assume that the concentration is right for a dust explosion. The particulate concentration levels within an operating fluidized bed will exceed the upper explosive limit so that an ignition starting below the surface of the bed is virtually ruled out. The concentration in the freeboard (above the bed) may be in the explosive range; however, the vigorous flow of air or gas through the bed will tend to quench any ignition before it really gets started. In manufacturing processes involving fluidization, the most serious powder explosion risk is probably going to be associated with filling or emptying operations, or when the bed is started up or shut down. In all these situations, there will exist a transient period when the dust concentration in the vessel—probably in the freeboard region—will be in the explosive range. Numerous occurrences of explosions during filling and emptying of storage silos and hoppers have been reported, and there is every reason to anticipate similar hazards in the filling and emptying of fluidized bed systems. Therefore, measures taken to avoid risks in hoppers and silos are appropriate for fluidized beds. These measures are discussed in Secs. 5.1 and 5.2. Minimum Ignition Energy of Powders. With dust and oxygen concentrations in the correct range, there is still the requirement of a ignition source before a fire or explosion can occur. The more well-known ignition sources—open flames, overheated bearings, and electrical sparks caused by short-circuits—are examples where the available energy is adequate for igniting virtually any suspended dust. On the other hand, electrostatic sparks usually have a limited pool of available energy and, quite often, the electrostatic energy released in an electrostatic discharge is only just comparable to the ignition requirement of the dust. Given the uncertainty of the estimates for the parameters used, the assessment of ESD hazards is quite problematic. Usually, close attention is paid to the minimum ignition energy (MIE) of the suspended dust.

840 Fluidization, Solids Handling, and Processing Table 5. Representative ignition temperatures, minimum explosion concentration, and minimum ignition energy for selected dusts, from NFPA Fire Protection Handbook (National Fire Protection Association, 1986) and other sources.

Types of dust

Ignition temp. of dust cloud, °C

Ignition temp. of dust layer, °C

Min. ignition energy (MIE), J

Min. explosion conc., g/m3

corn

400

250

0.04

55.

rice

510

450

0.10

85.

wheat flour

440

440

0.06

50.

wheat starch

430

-

0.025

50.

charcoal, wood

530

180

0.02

140.

KY coal (bit.)

610

180

0.03

50.

PA coal (anth.)

730

0.10

65.

CA lignite

450

− 200

0.03

30.

aspirin vitamin B1

660

−

0.025

50.

360

0.06

35.

vitamin C

460

− 280

0.06

70.

aluminum

610

326

0.01

45.

titanium

330

510

0.025

45.

uranium

20

100

0.045

60.

cell. acetate Nylon®

430

0.03

40.

500

− 430

0.02

30.

PMMA

480

−

0.02

30.

polycarbonate

710

−

0.025

25.

PPO

540

−

Agricultural

Carbonaceous

Drugs

Metals

Plastic resins, etc.

~0.05

~60.

Electrostatics and Dust Explosions 841 To measure the MIE of a powder, a sample of the material is placed in a vessel called a Hartmann cell, the powder is dispersed by a strong jet of air, and then a capacitive spark of well-characterized energy and time duration is initiated between two electrodes mounted inside the vessel. The spark energy is increased incrementally under ignition occurs. Despite standardized procedures established for these measurements (British Standards Institute, 1991; ISC, 1994), reliable measurement of the minimum ignition energy in dusts is notoriously difficult. There is of course the problem of dispersing the dust reproducibly and uniformly, but there also exists controversy about the nature of the igniting spark. Investigators have found that the wave shape and duration of the current pulse associated with the spark—controlled by the series impedance of the discharge circuit— influence the measured value of the MIE (Eckhoff, 1975; Field, 1982). It is now suspected that the true MIE values of many dusts may be lower by factors from two to five than the accepted data of twenty years ago (Taylor and Secker, 1994). The MIE data in Table 5 must be questioned for this reason. In fact, a much more conservative value of 5 mJ has now been adopted for the MIE of manufactured polymer powders. Another important discovery is that the measured MIE of powders depends very strongly on particle size. For example, the measured values for polyethylene are 500 mJ and 10 mJ respectively for narrow size cuts centered at 100 µm and 30 µm. This result indicates how strongly fines can influence explosion hazards in powders. In fact, Britton argues that it is the surface-average diameter that provides the best correlation to measured MIE (Britton, 1992). Uncertainties about the accuracy of measured MIE values not withstanding, the data provided in Table 5 do reveal that the dusts of a wide range of commercial and industrial materials—from agricultural grains to powdered metals and polymer plastics—pose an undeniable explosion or fire hazard when dispersed in air.

4.2

Types of Discharges

An electrostatic dust ignition can occur when the discharge releases sufficient thermal energy within a sufficiently short period of time and small volume to ignite the suspended dust. Electrostatic ignition is complicated by the fact that there are a number of distinct ESD mechanisms important in electrostatic hazards and hazard abatement (Glor, 1988). Capacitive Discharges. Capacitive discharges are responsible for at least 90% of all dust and vapor ignitions of ESD origin. The

842

Fluidization,

Solids Handling,

and Processing

requirements for these most well-known of discharges are two electrically isolated conducting bodies (usually metallic and one ofwhich is usually a grounded pipe or container wall). The metal object can be some ungrounded implement, such as a funnel or scoop shovel, or a metallic pipe or cable that has become ungrounded due to a broken electrical bond. There must exist sufficient potential difference and charge to cause an avalanche breakdown (spark) in the air between the conductors (refer to Fig. 5). The mechanism of charging can be conduction or induction and the charge can be positive or negative. Capacitivesparksreadily ignite many flammable dusts.

Figure 5. The capacitive discharge. Sometimes series resistance slows the dis. charge, increasing or decreasing the likelihood of a dust ignition.

Capacitive discharges are characterized by the electrostatic energy Ueconverted to thermal energy in the spark. While somewhat difficult to measure accurately, the energy released in a capacitive spark is commonly estimated using the following relationship. Eq. (10)

Ue = Y2CobjectJl2 = ~/2Cobject

Here, Cobjectis the capacitance and V the voltage between the two metal conductors. The capacitanceof typical chargeablecomponents dependson their size, shape, and distance from a grounded surface. For purposes of electrostatic hazard assessment, it is usually convenient to rely upon measured data like that in Table 6.

Electrostatics and Dust Explosions 843 Table 6. Representative Capacitance Values for Typical Components in a Plant or Commercial Facility Handling Powders (Eichel, 1967)

Item

Capacitance range (Cobject )

Small metal implements (funnels, scoops, etc.)

10 to 20 pF

Small metal containers (up to ~50 liters)

10 to 100 pF

Medium sized containers (up to ~200 liters)

50 to 300 pF

Miscellaneous plant components

100 to 1000 pF

Human body (depends on shoeware)

200 to 300 pF

Filter receiver components

10 to 100 pF

Typical large transport truck Lined cylindrical vessel (4 m diameter)

~1000 pF ~100,000 pF

Two distinct conditions must be met in order for a capacitive spark to ignite a flammable dust. First, the voltage difference between the object and ground must be sufficiently high to promote a discharge. The sparking potential is a complex function of the capacitor’s geometry and the length of the gap across which the discharge must jump. At standard atmospheric conditions, the minimum sparking potential is Vmin ≈ 350 V and it is achieved at a gap spacing of 6 to 7 µm. According to Paschen’s law, the sparking potential is higher for gaps larger or smaller than this value (Cobine, 1958). Figure 6 shows the dependence of the sparking potential upon the product of pressure and gap spacing. In general, the following condition on voltage is required for a spark to occur Eq. (11)

V > Vmin

In ordinary practice, the rather conservative value of 100 volts is recommended for Vmin to provide a margin of safety (Gibson, 1979). The second condition for an ignition is that the energy released in the spark Ue must exceed the minimum ignition energy (or MIE) of the dust. Eq. (12)

Ue > MIE


Sparking potential, V s (V)

1500

air at 20o C 1000

minimum sparking potential

500

0 0.0

5.0

10.0

15.0

20.0

25.0

pressure times spacing, pd (torr-mm) Figure 6. Sparking potential in air as a function of pd, the product of pressure and gap spacing. (Adapted from data of Cobine, 1958.)

Two other useful parameters related to capacitive sparks and dust ignitions are (i) the optimum sparking distance is ~10 mm and (ii) the quenching distance is ~7 mm. Refer to Sec. 4.1 and Table 5, where MIE is discussed in more detail. In a specific situation, determination of the voltage V for use in Eqs. (10) and (11) is based on having reliable estimates for charging current and leakage resistance (cf. Sec. II.B of Jones and King, 1991). Figure 7 contains a convenient nomograph for assessment of capacitive discharge ignition risks. With estimates for the capacitance and voltage of an object, the value of Ue is obtained by drawing a straight line between the points on these two scales and then reading the intercept with the middle scale. For convenience, the MIE values of some important powders are indicated on this scale, including the generally accepted value of 5 mJ for polymer powders. Note that, in general, MIE values for vapors are much lower.

Electrostatics and Dust Explosions

845

Capacitance, c (picofarads)

Figure 7. Nomograph for calculation of capacitive electrostatic discharge energies. {Bodurtha, 1980; Jones and King, 1991.)

846 Fluidization, Solids Handling, and Processing Brush Discharges. Brush discharges, which have effective energies in the range of ~1 to ~3 mJ, can occur between a conductor with radius of curvature in the range 5 to 50 mm (0.2" to 2") and either another conductor or a charged insulating surface when the field intensity at the surface of the curved conductor exceeds approximately 5 kV/cm (Maurer, 1979). This type of discharge, which differs from a capacitive spark in that it is more spatially diffuse, is regarded as incapable of igniting any but the most sensitive dusts (though it can ignite many vapors and gases). Brush discharges are favored over capacitive sparks when current flow is limited by the high electrical resistance of the charged insulating surface. Corona Discharges. Corona discharges, which occur when an electrode or any conductive object with radius of curvature less than ~5 mm experiences a strong electric field, are weaker and still more spatially diffuse than brush discharges. They are not sufficiently energetic to ignite even gases or vapors (Glor, 1988), so they present no direct ESD ignition hazard for suspended dusts; however, corona may contribute to accelerated charging of ungrounded metal objects or components within a vessel which may then lead to a capacitive discharge capable of igniting dust. When the objective is to avoid corona discharges, sharp edges and protrusions, such as threaded bolts, rough welding seams, etc., inside vessels should be eliminated. Corona-induced product degradation, while never documented, may well be a factor to consider in the manufacture of clear or light-colored plastics for products where appearance is important. On the other hand, corona discharges, promoted by internally mounted wires or sharp metal protrusions inside a vessel, have been proposed as a way to promote charge neutralization of powders entering a vessel. Figures 8(a) and (b) show two distinct schemes for discharging powder during filling operations (Blythe and Reddish, 1979; van de Weerd, 1974). While these schemes have been shown to be effective, there has been hesitancy to install them widely, a concern probably due to understandable worries about breakage. Propagating Brush Discharges. Conditions for a propagating brush discharge occur when a thin insulating layer, backed by a conductor, becomes charged (Glor, 1988). See Fig. 9(a). With spark energies up to ~1 J, these are the most energetic of all electrostatic discharges and they are capable of igniting virtually any flammable dust. Many serious ESD ignitions causing personal injuries as well as extensive plant damage have been attributed to propagating brush discharges (Hughes et al., 1975; Lüttgens, 1985). The high available energy for the propagating brush

Electrostatics and Dust Explosions 847 discharge comes from the thin insulating layer, which acts like a large distributed capacitor. Once a discharge is initiated, electrostatic energy over a considerable area can be converted to heat due to discharges flashing along the charged surface. For this reason, it is best to avoid nonconductive liners on the inside surfaces of vessels, bins, and pipes. Furthermore, it is vitally important not to use pipe, tubing, inserts, or any other components made of plastic in any place exposed to dust. For example, plastic pipe can become conductive on its outside surface due to moisture or contamination, providing the condition for a propagating brush discharge even though the pipe initially does not have a conductive backing (Lüttgens, 1985). Refer to Fig. 9(b) which shows a common and dangerous situation.

inlet pipe

powder flow

passive discharge wire

pointed discharge rod

plastic weight

powder heap

interior of vessel

(a)

(b)

Figure 8. Application of corona to reduce powder charge level in granular solids. (a) Suspended wire for neutralization of charge entering silo (Blythe and Reddish, 1979). (b) Inlet pipe equipped with grounded corona discharge point (van de Weerd, 1974). Corona discharge wires, rods and/or points should have a diameter from 1 to 3 mm and must be securely grounded.

848

Fluidization,

Solids Handling,

/ i. d

and Processing

surfacecharge

fluidization solids handling and processing

Fluidization, Solids Handling, and Processing: Industrial Applications

Powders and solids: developments in handling and processing technologies

Bulk solids handling: equipment selection and operation

Applications of Fluidization in Food Processing

Fluidization Dynamics

Guidelines for Safe Handling of Powders and Bulk Solids

Handbook of Conveying and Handling of Particulate Solids

Guidelines for Safe Handling of Powders and Bulk Solids

Fluidization Dynamics

Fluidization Dynamics

handbook of fluidization and fluid-particle systems

Livestock Handling and Transport

Handling Sin

Handbook of fluidization and fluid-particle systems

Dog Handling

Dog Handling

Handling Sin

Fluidization Dynamics: A Predictive Theory

Fluidization dynamics: a predictive theory

Preventing and handling product liability

Ship Handling Theory and Practice

Mechanics of solids and materials

Mechanics of Solids and Materials

Diffusion and Reactivity of Solids

Care and handling of CDs and DVDs

Fruit and vegetables: harvesting, handling, and storage

Mechanics of Solids and Materials

Constraint Handling Rules

Handling experimental data

Handling Complaints (Management Pocketbooks)

fluidization solids handling and processing